Integrated Process And Methods Of Producing (E)-1-Chloro-3,3,3-Trifluoropropene

ABSTRACT

The present invention relates to methods, process, and integrated systems for economically producing (E)-1-chloro-3,3,3-trifluoropropene via vapor phase and/or liquid processes.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is related to and claims the priority benefit of U.S.provisional application No. 61/305,803 filed Feb. 18, 2010 and U.S.provisional application No. 61/379,633 filed Sep. 2, 2010, the contentsboth of which are incorporated herein by reference.

FIELD OF THE INVENTION

This invention relates to processes, methods and systems for producinghydrochlorofluorooleins, particularly (E)1-chloro-3,3,3-trifluoropropene.

BACKGROUND OF THE INVENTION

Chlorofluorocarbon (CFC) based chemicals have been widely use inindustry in a variety of different applications including asrefrigerants, aerosol propellants, blowing agents and solvents, amongothers. Certain CFCs, however, are suspected of depleting the Earth'sozone layer. Accordingly, more environmentally friendly substitutes havebeen introduced as replacements for CFCs. One example of such asubstitute is 1,1,1,3,3-pentafluoropropane (HFC-245fa). HFC-245fa isrecognized as having favorable physical properties for certainindustrial applications, such as foam blowing agents and solvents, andtherefore is considered to be a good substitute for the CFCs previouslyused for these applications. Unfortunately, the use of certainhydrofluorocarbons, including HFC-245fa, in industrial applications isnow believed to contribute to global warming. As a result, moreenvironmentally friendly substitutes for hydrofluorocarbons are nowbeing sought.

The compound I-chloro-3,3,3-trifluoropropene, also know as HCFO-1233zdor simply 1233zd, is a leading candidate for replacing HFC-245fa in someapplications, including blowing agents and solvents. 1233zd has aZ-isomer and an E-isomer. Due to differences in the physical propertiesbetween these two diastereoisomers, pure 1233zd(E), pure 1233zd(Z), orcertain mixtures of the two isomers may be suitable for particularapplications as refrigerants, propellants, blowing agents, solvents, orfor other uses.

Processes for synthesizing 1233zd are known. WO 97/24307, for example,discloses a process for preparing 1233zd via the gas-phase reaction of1,1,1,3,3-pentachloropropane (HCC-240fa) with hydrogen fluoride (HF).This process, however, produces relatively low yields of 1233zd. U.S.Pat. No. 6,844,475 describes a liquid phase reaction of HCC-240fa withHF to produce 1233zd in higher yields. A preferred temperature range forthis reaction was purported to be about 50° C. to about 120° C., withspecific examples being demonstrated at 90° C. (resulting in a 1233zdyield of about 80 wt. %) and 120° C. (resulting in a 1233zd yield ofover 90 wt. %). The 1233zd yield at the lower temperature is notparticularly good. The yield at the higher temperature is good, butApplicants have found that operating the process at this temperature andabove produces increased amounts of the Z-isomer. Accordingly, thereremains a need for a process for selectively producing 1233zd(E) in highyields. This application satisfies that need among others.

SUMMARY OF THE INVENTION

Applicants have unexpectedly found that a 1233zd reaction product havinga majority of E-isomer is formed by carefully maintaining thetemperature of either (1) a catalytic liquid phase reaction between HFand one or a combination of HCC-240fa, 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene or (2) a catalytic vapor phase reactionbetween HCC-240fa and HF. With respect to the former, the temperature ismaintained preferably, though not exclusively, at about 85° C. to about120° C. Applicants have also discovered that the E-isomer yield isfurther improved by passing the liquid phase product through a strippingcolumn as it leaves the reactor to reflux unreacted reactants back tothe reactor, particularly, though not exclusively, when the strippingcolumn is operated to achieve an average stripping temperature of about10° C. to about 40° C. below the corresponding reaction temperature.With respect to the gas phase reaction, the temperature is maintainedpreferably, though not exclusively, at about 200 to about 450° C., witha pressure being maintained at about 0 to about 160 psig.

Overall process efficiency is achieved by integrating other unitoperations to separate and recycle unreacted reactants and/orundesirable isomers of 1233zd, and also to separate and removeby-products. Thus, in certain preferred embodiments, the process isdirected to an integrated process for producing 1233zd(E) in highyields.

In one aspect, the instant invention relates to a method or process forproducing a chlorofluoroalkene comprising: (a) providing a liquidreaction admixture comprising hydrogen fluoride,1,1,1,3,3-pentachloropropane (or a hydrohalocarbon mixture of1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene), and a fluorinated metal chloride catalyst,wherein said hydrogen fluoride and 1,1,1,3,3-pentachloropropane (or thehydrohalocarbon mixture) are present in a molar ratio of greater thanabout 3:1 and wherein said a fluorinated metal chloride catalyst isselected from the group consisting of partially or fully fluorinatedTiCl₄, SnCl₄, TaCl₅, SbCl₃, FeCl₃, or AlCl₃; and (b) reacting saidhydrogen fluoride and 1,1,1,3,3-pentachloropropane (or thehydrohalocarbon mixture of 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene) in thepresence of said catalyst in a liquid phase and at a reactiontemperature of about 85° C. to about 120° C. to produce a reactionproduct stream comprising (E)1-chloro-3,3,3-trifluoropropene, hydrogenchloride, unreacted hydrogen fluoride, entrained catalyst,(Z)1-chloro-3,3,3-trifluoropropene and, optionally, unreactedhydrohalocarbon starting product (e.g. 1,1,1,3,3-pentachloropropane,and/or 1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene),wherein said product stream has a weight ratio of(E)1-chloro-3,3,3-trifluoropropene to (Z)1-chloro-3,3,3-trifluoropropeneof greater than 1.

This method or process may further include the step of (c) contactingsaid reaction product stream with a heat exchanger to produce (i) afirst crude product stream comprising a majority of said hydrogenchloride, a majority of said (E)1-chloro-3,3,3-trifluoropropene,optionally a majority of said (Z)1-chloro-3,3,3-trifluoropropene, and atleast a portion of said unreacted hydrogen fluoride, wherein saidportion is an amount sufficient to form an azeotrope with one or more ofsaid (E)1-chloro-3,3,3-trifluoropropene and said(Z)1-chloro-3,3,3-trifluoropropene, and (ii) a reflux componentcomprising a majority of said entrained catalyst and said unreactedhydrogen fluoride; and (d) returning said reflux component to saidreaction admixture.

In further embodiments, the method or process further includes one ormore of the following steps: (e) separating unreacted reactants,including unreacted 1,1,1,3,3-pentachloropropane, and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene, viadistillation and recycling these unreacted reactants back to thereactor; (f) removing at least a portion, and preferably a majority, ofhydrochloric acid by-product; (g) separating and recycling unreacted HFin a crude product stream via a sulfuric acid adsorption or a phaseseparation; (h) distillation of the crude product stream to separate(E)1233zd from reaction by-products; and (i) isomerization of 1233zd(Z)by-products to form 1233zd(E).

The foregoing steps may be provided as an integrated system forproducing a hydrofluoroolefin comprising (a) one or more feed streamscumulatively comprising hydrogen fluoride and1,1,1,3,3-pentachloropropane (or a hydrohalocarbon mixture of1,1,1,3,3-pentachloropropane, and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene); (b) a liquid phase reactor charged with aliquid phase fluorination catalyst and maintained at a first temperatureof about 85° C. to about 120° C., wherein said liquid phase reactor isfluidly connected to said one or more feed streams; (c) a strippingsystem comprising a stripping column having an average temperaturemaintained at a second temperature of about 10° C. to about 40° C. belowsaid first temperature, a reflux stream fluidly connected to saidstripping column, and a first crude product stream fluidly connected tosaid stripping column, wherein said reflux stream is fluidly connectedto said liquid phase reactor; (d) a hydrogen chloride removal systemcomprising a first distillation column, a hydrogen chloride by-productstream fluidly connected to said first distillation column, and a secondcrude product stream fluidly connected to said first distillationcolumn, wherein said first distillation column is fluidly connected tosaid stripping column; (e) a hydrogen fluoride recovery systemcomprising a sulfuric acid absorption and recycle system or a phaseseparation vessel, a second recycle stream comprising hydrogen fluoridefluidly connected to said sulfuric acid absorption and recycle system ora phase separation vessel, a third product stream comprising (E) and (Z)1-chloro-3,3,3-trifluoropropene fluidly connected to said sulfuric acidabsorption and recycle system or a phase separation vessel, wherein saidsulfuric acid absorption and recycle system or a phase separation vesselis fluidly connected to said second crude product stream; and (f) a1-chloro-3,3,3-trifluoropropene purification system comprising a seconddistillation column fluidly connected to said third product stream; afinal product stream comprising (E)1-chloro-3,3,3-trifluoropropenefluidly connected to said second distillation column; a secondby-product stream fluidly connected to said distillation column, aisomerization reactor fluidly connected to said second by-productstream; an a product recycle stream fluidly connected to saidisomerization reactor and said second distillation column.

In another embodiment, the instant invention relates to a method orprocess for preparing (E)1-chloro-3,3,3-trifluoropropene by firstproviding 1,1,1,3,3-pentachloropropane with hydrogen fluoride in a vaporphase reaction mixture and in the presence of a fluorinated catalyst.The mixture is provided within a reactor where hydrogen fluoride and1,1,1,3,3-pentachloropropane are present in a HF:organic molar ratio ofgreater than about 3:1. 1,1,1,3,3-pentachloropropane is reacted withhydrogen fluoride in the presence of the catalyst and at a reactiontemperature of about 200 to about 450° C. and a pressure of about 0 toabout 160 psig. The resulting product stream includes(E)1-chloro-3,3,3-trifluoropropene, hydrogen chloride, unreactedhydrogen fluoride, unreacted 1,1,1,3,3-pentachloropropane, reactionby-products, and optionally (Z)1-chloro-3,3,3-trifluoropropene.

The fluorinated catalyst for the vapor phase reaction may be selectedfrom one or more catalysts in the group of chromium based catalysts,aluminum based catalysts, cobalt based catalysts, manganese basedcatalysts, nickel and iron oxide based catalysts, hydroxide basedcatalysts, halide based catalysts, oxyhalide based catalysts, inorganicsalts thereof or mixtures thereof. In one embodiment, the catalyst isselected from Cr₂O₃, Cr₂O₃/Al₂O₃, Cr₂O₃/AlF₃, Cr₂O₃/carbon,CoCl₂/Cr₂O₃/Al₂O₃, NiCl₂/Cr₂O₃/Al₂O₃, CoCl₂/AlF₃, NiCl₂/AlF₃ andmixtures thereof. In other embodiments, the catalyst is selected fromFeCl₃/C, SnCl₄/C, TaCl₅/C, SbCl₃/C, AlCl₃/C, and AlF₃/C.

The product or final reaction stream in the vapor phase reaction mayundergo post-reaction processing to isolate the desired products andrecycle certain unreacted starting reagents. In one embodiment, such aprocess step includes separating the reaction product stream to produce(a) a first overhead stream of HCl, and (b) a first bottoms streamincluding unreacted hydrogen fluoride, unreacted1,1,1,3,3-pentachloropropane, (E)1-chloro-3,3,3-trifluoropropene, andoptionally (Z)1-chloro-3,3,3-trifluoropropene. The first bottoms streammay then be separated to produce (a) a recycle stream of HF, which maybe recycled back to the vapor phase reaction; and (b) a stream thatincludes unreacted 1,1,1,3,3-pentachloropropane, reaction by-products,(E)1-chloro-3,3,3-trifluoropropene, and optionally(Z)1-chloro-3,3,3-trifluoropropene. This stream can be then separatedfurther to produce (a) a stream of unreacted1,1,1,3,3-pentachloropropane and reaction by-products; and (b) a streamcomprising (E)1-chloro-3,3,3-trifluoropropene, and optionally(Z)1-chloro-3,3,3-trifluoropropene. In embodiments where both the E andZ isomer are provided, the components of said stream are then separated,i.e., (E)1-chloro-3,3,3-trifluoropropene is separated from(Z)1-chloro-3,3,3-trifluoropropene. The(Z)1-chloro-3,3,3-trifluoropropene may then be isomerized to produce(E)1-chloro-3,3,3-trifluoropropene. In any of the foregoing steps theseparation technique includes, but is not limited to, distillation oradsorption, but may also be adapted to use alternative separationtechniques that are known in the art.

The foregoing steps for the vapor phase reaction may be provided in anintegrated system for producing a hydrofluoroolefin including (a) one ormore feed streams cumulatively of hydrogen fluoride and1,1,1,3,3-pentachloropropane; (b) a vapor phase reactor with a vaporphase fluorination catalyst and maintained at a first temperature ofabout 200 to about 450° C. and a pressure of about 0 to about 160 psig,wherein said vapor phase reactor is fluidly connected to said one ormore feed streams; (c) a hydrogen chloride removal system including afirst distillation column, a hydrogen chloride by-product stream fluidlyconnected to said first distillation column, and a crude product streamfluidly connected to said first distillation column wherein said firstdistillation column is fluidly connected to said vapor phase reactor;(d) a hydrogen fluoride recovery system including a sulfuric acidstripping and recycle system or a phase separation vessel, a recyclestream comprising hydrogen fluoride fluidly connected to said sulfuricacid stripping and recycle system or a phase separation vessel, aproduct stream including (E) and (Z) 1-chloro-3,3,3-trifluoropropenefluidly connected to said sulfuric acid stripping and recycle system ora phase separation vessel, wherein said sulfuric acid stripping andrecycle system or a phase separation vessel is fluidly connected to saidcrude product stream; (e) a 1-chloro-3,3,3-trifluoropropene purificationsystem including a second distillation column fluidly connected to saidproduct stream; a final product stream including (E)1-chloro-3,3,3-trifluoropropene fluidly connected to said seconddistillation column; a second by-product stream fluidly connected tosaid distillation column, a isomerization reactor fluidly connected tosaid second by-product stream; and a product recycle stream fluidlyconnected to said isomerization reactor and said second distillationcolumn.

Additional embodiments and advantages to the instant invention will bereadily apparent to one of skill in the art based on the disclosureprovided below.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 shows a schematic depiction of an integrated liquid phasesynthesis of 1233zd (E) according to a preferred embodiment of theinvention;

FIG. 2 shows a schematic depiction of an integrated liquid phasesynthesis of 1233zd(E) according to another preferred embodiment of theinvention;

FIG. 3 shows a schematic depiction of an integrated liquid phasesynthesis of 1233zd(E) according to another preferred embodiment of theinvention; and

FIG. 4 shows a schematic depiction of an integrated liquid phasesynthesis of 1233zd(E) according to yet another preferred embodiment ofthe invention.

FIG. 5 is a schematic diagram of an example of a vapor phase system thatproduces (E)1-chloro-3,3,3-trifluoropropene.

FIG. 6 is another schematic diagram of another vapor phase system thatproduces (E)1-chloro-3,3,3-trifluoropropene.

FIG. 7 is a further schematic diagram of a further vapor phase systemthat produces (E)1-chloro-3,3,3-trifluoropropene.

FIG. 8 is yet another schematic diagram of yet another vapor phasesystem that produces (E)1-chloro-3,3,3-trifluoropropene.

DETAILED DESCRIPTION OF PREFERRED EMBODIMENTS OF THE INVENTION

In certain embodiments, the present invention relates to methods andprocess for continuously and economically producing(E)-1-chloro-3,3,3-trifluoropropene via a fully integrated liquid phaseor vapor phase process. Representative fully integrated processes formaking (E)-1-chloro-3,3,3-trifluoropropene are described below.

Liquid Phase Reaction

The reaction chemistry for the liquid phase process involves asingle-step reaction of 1,1,1,3,3-pentachloropropane (or ahydrohalocarbon mixture of 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene) withanhydrous HF in a liquid-phase, catalyzed reactor to produce primarily(E)-1-chloro-3,3,3-trifluoropropene (1233zd(E)) plus HCl as aby-product. Preferably, the reaction is maintained under conditions(temperature, pressure, residence time) to increase the relative ratioof (E) to (Z) isomers of 1233zd while also minimizing the reaction of HFwith the resulting 1233zd(E) which would lead to the formation ofHFC-244fa, which in turn can react further to produce HFO-1234ze.Accordingly, the desired reactions involve:

Undesired reactions, which are preferably avoided, include:

In certain embodiments, the manufacturing process for the liquid phasereaction comprises six major unit operations: (1) catalyst preparation(preferably titanium tetrachloride), (2) fluorination reaction(continuous or semi-batch mode) using HF with simultaneous removal ofby-product HCl and the product 1233zd(E), (3) separation andpurification of by-product HCl, (4) separation of excess HF back to (2),(5) purification of final product, 1233zd(E), and (6) isomerization ofby-product 1233zd(Z) to 1233zd(E) to maximize the process yield. Therelative positions of these operations are shown in FIGS. 1-4.

Unit Operation One: Catalyst Preparation

In one aspect, the fluorination reaction described herein uses a liquidphase catalyst of proper strength to achieve the desired reactionpreferentially. One fluorination catalyst used is titanium tetrachloride(liquid under ambient conditions) which has been partially or totallyfluorinated by the action of anhydrous HF. This catalyst unexpectedlyachieves the desired degree of conversion without forming a significantamount of undesired volatile by-products (although formation of amoderate amount of HCl is unavoidable). The catalyst fluorination isconducted by adding a specified amount of HF and fluorination catalystto a reaction vessel equipped with an agitator at a temperature in therange of 0-120° C. Optionally, catalyst fluorination can be conducted bycombining HF, fluorination catalyst, and 1,1,1,3,3-pentachloropropane or1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene mixture in the reactor provided the reactortemperature is below the 1,1,1,3,3-pentachloropropane or1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene mixture fluorination reaction initiationtemperature (<85° C.) for safety considerations. Order of fluorinationcatalyst, HF, and/or 1,1,1,3,3-pentachloropropane or1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene mixture addition to reactor is not important,but addition of HF first is preferred. Partial or complete fluorinationof the catalyst will occur upon addition and HCl by-product will begenerated increasing the pressure within the reaction vessel which isthen controlled at the desired reaction pressure. Additionalfluorination catalysts that can be used include SnCl₄, TaCl₅, SbCl₃,FeCl₃, and AlCl₃ which have been partially or totally fluorinated by theaction of anhydrous HF. Preferably the reactor is constructed frommaterials which are resistant to the corrosive effects of the HF andcatalyst, such as Hastelloy-C, Inconel, Monel, Incalloy, orfluoropolymer-lined steel vessels. Such liquid-phase fluorinationreactors are well known in the art.

Unit Operation Two: Fluorination Reaction and Reactor and StrippingColumn

The arrangement and operation of the reactor and stripping column isparticularly important in achieving a high yield of 1233zd(E). In apreferred embodiment, the reaction is conducted of an agitated,temperature-controlled reactor containing the liquid fluorinationcatalyst. One or more feeds comprising hydrogen fluoride and1,1,1,3,3-pentachloropropane or 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene mixtureenter the reactor where they contact each other and the catalyst in aliquid phase. The resulting reaction produces a gas phase productcomprising 1233zd(E) as well as various other by-products including HCland possibly 1233zd(Z). The gas phase product leaves the liquid phasereactor and enters an integrated distillation column (operating instripping mode) which permits the desired product to leave (along withby-product HCl, traces of light organics [principally 1234ze(E+Z)], andsufficient anhydrous hydrogen fluoride (AHF) to form the azeotropes),while retaining the bulk of the HF, plus under-fluorinated and dimerizedorganics, plus fluorination catalyst entrained in the gas stream. Oncethe catalyst has been prepared, the reaction can be initiatedimmediately upon heating to the desired reaction temperature. The flowof HF needed for the catalyst preparation can be resumed, and additionof the 1,1,1,3,3-pentachloropropane or 1,1,1,3,3-pentachloropropaneand/or 1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropenemixture can be started immediately to cause continuous reaction.Alternatively, a large amount of the same 1,1,1,3,3-pentachloropropaneor 1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene mixture can be added at one time as a batchcharge, and then HF can be added gradually to the reactor (a semi-batchoperation). Alternatively, a large amount of HF can be added at one timeas a batch charge, and then the same 1,1,1,3,3-pentachloropropane or1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene mixture can be added gradually to the reactor(a semi-batch operation). Proper temperature control of the coolant andsufficient reflux action are desirable for optimum operation of thestripping column to be effective. General operating conditions which wehave found to work well for the reaction and stripping are: Operatingpressure of 80-140 psig maintained by a control valve on the exitingflow from the stripper column; reactor temperature of 85-120° C.,primarily supplied by steam flow into the reactor jacket; application of−40° C.-25° C. brine cooling to the heat exchanger on top of thestripper column to induce reflux; temperature in the center portion ofthe stripper about 10-40° C. below that in the reactor; additional heatinput by superheating the HF vapor feed with high-pressure steam to120-150° C.; feed rate of HF to maintain reactor and stripperconditions, typically 0.5 to 2.0 pounds per hour in apparatus of thissize.

It has been discovered that maintaining the reaction under the operatingconditions, particularly, a temperature range of 85-120° C., morepreferably 90-110° C., and most preferably 95-100° C., produces anunexpected shift in the reaction mechanism which produces a high ratioof 1233zd(E) compared to 1233zd(Z).

Unit Operation Three: Removal of HCl

The HCl formed continuously during the reaction is removed from thereactor due to its volatile nature, and flows through the attacheddistillation column without condensing. The material can then bepurified and collected for sale (or further purification) by using alow-temperature HCl distillation column. High purity HCl is isolated andcan be absorbed in de-ionized water as concentrated HCl for sale.

Unit Operation Four: Separation and Recycle of Excess HF Back to UnitOperation Two

The bottoms stream from the HCl removal column (unit operation three)contains a crude product mixture of 1233zd(E) and HF (in someembodiments about 30 wt %) is fed to a sulfuric extractor or a phaseseparator for removal of HF from this mixture. HF is either dissolved inthe sulfuric acid or phase separated from the organic mixture. Forembodiments utilizing a sulfuric acid adsorption system, the HF is thendesorbed from the sulfuric acid/HF mixture by stripping distillation andrecycled back to the reactor. For embodiments utilizing a phaseseparator, HF is phase-separated and recycled back to the reactor. Theorganic mixture either from the overhead of the sulfuric acid extractoror from the bottom layer of the phase separator may require treatment(scrubbing or adsorption) to remove traces of HF before it is fed to thenext unit operation.

Unit Operation Five: Purification of Final Product

Purification of final product preferably comprises two continuouslyoperating distillation columns. The 1^(st) column is used to removelight ends from the 1233zd(E) and the 2^(nd) column is used to removethe heavier components, primarily the 1233zd(Z), which is fed to anisomerization reactor, collected for further use or optionally recycledback to the reactor (i.e., unit operation two). In certain embodiments,it is desirable to have a purge of heavy by-products from this stream.

Unit Operation Six: Isomerization of by-Product 1233zd(Z) to 1233zd(E)

To maximize the 1233zd(E) yield in this process, the by-product1233zd(Z) formed in the reaction and exiting the bottom of the 2^(nd)column is fed as a vapor to a reactor that contains an isomerizationcatalyst, preferably fluorinated chromium oxide. Here, the by-product isconverted to the desired product. The isomerization reactor exit streamis then recycled to unit operation four for purification.

In certain preferred embodiments, this step involves controlling thetemperature of a heated surface to greater than 50° C.-350° C. Theheated surface is contacted with the stream containing the 1233zd(Z)by-product. The feed stream is contacted with the heated surface for aperiod of time sufficient to convert at least a portion of the 1233zd(Z)to 1233zd(E) to produce a product stream rich in 1233zd(E).

In some embodiments, the heated surface includes the inside of a reactorvessel. In addition, or in the alternative, the heated surface mayinclude an outer surface of a packing material, for example a packingmaterial that is packed in a reaction vessel. In some embodiments, thereactor vessel is a batch-wise reactor vessel that can be charged withthe feed stream. In some such embodiments, the feed stream may be sealedin the batch-wise reactor, and, after sufficient time passes toisomerize the desired amount of 1233zd(Z), the reactor vessel may beopened to remove the product stream. In other embodiments, the reactorvessel is a continuous-type reactor vessel, for example a reactor vesselwith a first opening and a second opening and a fluid pathway betweenthe first and second openings. The feed stream is fed into the reactorvessel through the first opening and passes through the reactor vesselat a rate sufficient to isomerize the desired amount of 1233zd(Z). Theresulting product stream exits the second opening. In one example, thereactor vessel is an elongate reactor vessel (e.g., a MONEL™ tube) withthe first opening at a first end and the second opening at a second end.

In some embodiments, the reactor vessel may be partially or entirelypacked with packing material, for example with a stainless steelpacking. In some embodiments, the relatively large surface area of thepacking material may facilitate the conversion reaction from the (Z) tothe (E) isomer. Support structures that support the packing material mayalso be disposed in or on the reactor vessel. For example, the packingmaterial may be supported by a mesh or other structure that is disposedunder, around, and/or within the packing material. The support structuremay comprise the same material as the packing material (e.g., stainlesssteel), nickel, or any other suitable material.

The packing materials may also comprise one or more catalyst materials.Examples of suitable catalysts for the isomerization of 1233zd are metaloxides, halogenated metal oxides, Lewis acid metal halides, zero-valentmetals, as well as combinations of these catalysts. Specific examples ofsuitable catalysts are AlF₃, Cr₂O₃, fluorinated Cr₂O₃, zirconium oxideand halogenated versions thereof, or an aluminum oxide and halogenatedversions thereof. In addition, the catalysts may be activated prior touse. Examples of activation procedures for several suitable catalystsmay be found in U.S. Publication No. 2008-0103342, which is herebyincorporated by reference in its entirety.

Turning to the Figures, FIG. 1 shows the synthesis of 1233zd(E) via aliquid phase reaction integrated process having a sulfuric acid HFrecovery system. Here, liquid phase reactor R1 is first charged with anamount of anhydrous hydrogen fluoride that is in a stoichiometric excessof the amount needed to totally fluorinate a metal chloride catalystfluorination catalyst, e.g. when using TiCl₄ a>4:1 mole ratio of HF tocatalyst is added. This is followed by the addition of a fluorinationcatalyst alone or in combination from the group comprising TiCl₄, SnCl₄,TaCl₅, SbCl₃, FeCl₃, or AlCl₃ to prepare the catalyst. TiCl₄ is mostpreferred. The catalyst preparation is done while the reactor is at0-120° C. HCl is generated during catalyst preparation and can be ventedout of the top of the catalyst stripper column CS-1 to control thereactor pressure at or below the intended operating pressure of thereactor. Preferably the reactor is constructed from materials which areresistant to the corrosive effects of the HF and catalyst, such asHastelloy-C, Inconel, Monel, Incolloy, or fluoropolymer-lined steelvessels. Such liquid-phase fluorination reactors are well known in theart. Additional HF and 1,1,1,3,3-pentachloropropane or1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene mixture is then added until good agitation isachieved.

The reaction mixture is then heated to about 85° C. where thefluorination reaction between HCC-240fa and HF is initiated. Continuous1,1,1,3,3-pentachloropropane or 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene mixture andHF (in a stoichiometric excess) feeds are simultaneously fed to heaterHX-1 and then into a liquid phase reactor R-1. Optionally,1,1,1,3,3-pentachloropropane or 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene mixture isfed directly into reactor R-1 and not through heater HX-1. The operatingpressure of 60-160 psig (preferably 80-140 psig) is maintained by acontrol valve on the exiting flow from the catalyst stripper column CS-1and the reactor temperature is kept in the range of 85-120° C.,primarily supplied by steam flow into the reactor jacket. A catalyststripper column CS-1 is connected to the reactor, R-1, and serves thepurpose of knocking down and returning entrained catalyst, some HF,partially fluorinated intermediates, and some unreacted1,1,1,3,3-pentachloropropane, and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene back to the reactor for further reaction.

The stream exiting the top of catalyst stripper CS-1 comprising mainly1233zd(E), 1233zd(Z), HF, and HCl (with some minor components includingpartially fluorinated intermediates and by-products, overfluorinatedby-products, and 1233zd dimers), then enters HCl column D-1. A streamcomprising mainly HCl by-product exits the top of the HCl column and isfed to an HCl recovery system. The recovered HCl by-product can be soldfor profit. The HCl column bottoms stream consisting mainly of1233zd(E), 1233zd(Z), and HF are then fed into an HF recovery system.The HF recovery system starts with the crude 1233zd/HF stream beingvaporized in heat exchanger HX-2 and fed into HF absorption column A-1.Here a liquid stream of 50-80% H₂SO₄ contacts the gaseous 1233zd/HFstream and absorbs the majority of the HF. The stream exiting the bottomof A-1 comprises HF/H₂SO₄/H₂O and is fed to heat exchanger HX-3 where itis heated to a temperature sufficient to flash the majority of the HFalong with small amounts of H₂O and H₂SO₄. This stream is fed to HFrecovery distillation column D-2. The liquid remaining after the HF isflashed off in HX-3 consisting mainly of H₂SO₄ and H₂O (with 0-2% HF) iscooled in HX-4 and recycled back to HF absorption column A-1. The HFrecovery column, D-2, bottoms stream comprising mainly H₂SO₄ and H₂O arerecycled back to heat exchanger HX-3. Anhydrous HF is recovered from thetop of the HF recovery column, D-2, and is recycled back to the reactorR-1 via vaporizer HX-1. The stream exiting the top of HF absorptioncolumn A-1 comprising mainly 1233zd(E) and 1233zd(Z) (trace HF) is sentforward to a polishing system A-2 where the gaseous stream contacts awater or a caustic solution to remove trace HF and is subsequently driedwith a desiccant. Acid free crude product exiting absorber A-2 is sentto the first of two purification columns, D-3. A stream exiting the topof the column D-3 consists mainly of reaction by-products that haveboiling points lower than that of 1233zd(E). The stream exiting thebottom of lights column D-3 consisting mainly of 1233zd(E) and 1233zd(Z)and heavier by-products is fed to product recovery distillation columnD-4. Product grade 1233zd(E) exits the top of the column to productstorage. The product column bottoms consist mainly of 1233zd(Z) andreaction by-products with boiling points higher than that ofHCFO-1233zd(E) is then fed to vaporizer HX-5 and then to isomerizationreactor R-2 where by-product 1233zd(Z) is converted to the desiredproduct. The stream leaving R-2 is then recycled to lights distillationcolumn D-3 for purification. Optionally, if any by-products in thestream entering R-2 are unstable they may decompose and form smallamounts of HF or HCl. In this case, the stream exiting R-2 can berecycled and combined with the stream entering the polishing system A-2to remove the acid. Optionally, the stream exiting the bottom of theproduct recovery distillation column, D-4 can be recycled back to liquidphase reactor R-1. In any of these options a heavies purge stream fromthe bottom of the product recovery distillation column, D-4, will berequired to prevent build-up of high boiling impurities in thepurification system. The heavies purge stream is collected for later useor waste disposal.

Referring to FIG. 2, shown is the synthesis of 1233zd(E) via a liquidphase reaction integrated process having sulfuric acid HF recovery andoptional recycle column after the reactor. Here, liquid phase reactor R1is first charged with an amount of anhydrous hydrogen fluoride that isin a stoichiometric excess of the amount needed to totally fluorinate ametal chloride catalyst fluorination catalyst, e.g. when using TiCl₄a>4:1 mole ratio of HF to catalyst is added. This is followed by theaddition of a fluorination catalyst alone or in combination from thegroup comprising TiCl₄, SnCl₄, TaCl₅, SbCl₃, FeCl₃, or AlCl₃ to preparethe catalyst. TiCl₄ is most preferred. The catalyst preparation is donewhile the reactor is at 0-120° C. HCl is generated during catalystpreparation and can be vented out of the top of the catalyst strippercolumn CS-1 to control the reactor pressure at or below the intendedoperating pressure of the reactor. Preferably the reactor is constructedfrom materials which are resistant to the corrosive effects of the HFand catalyst, such as Hastelloy-C, Inconel, Monel, Incolloy, orfluoropolymer-lined steel vessels. Such liquid-phase fluorinationreactors are well known in the art. Additional HF and1,1,1,3,3-pentachloropropane or 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene mixture isthen added until good agitation is achieved.

The reaction mixture is then heated to about 85° C. where thefluorination reaction between 1,1,1,3,3-pentachloropropane or1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene mixture and HF is initiated. Continuous1,1,1,3,3-pentachloropropane or 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene mixture andHF (in a stoichiometric excess) feeds are simultaneously fed to heaterHX-1 and then into a liquid phase reactor R-1. Optionally,1,1,1,3,3-pentachloropropane or 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene mixture isfed directly into reactor R-1 and not through heater HX-1. The operatingpressure of 60-160 psig (preferably 80-140 psig) is maintained by acontrol valve on the exiting flow from the catalyst stripper column CS-1and the reactor temperature is kept in the range of 85-120° C. primarilysupplied by steam flow into the reactor jacket. A catalyst strippercolumn CS-1 is connected to the reactor, R-1, and serves the purpose ofknocking down and returning entrained catalyst, some HF, partiallyfluorinated intermediates, and some unreacted1,1,1,3,3-pentachloropropane, 1,1,3,3-tetrachloropropene or1,3,3,3-tetrachloropropene back to the reactor for further reaction.

The stream exiting the top of catalyst stripper CS-1 comprising mainly1233zd(E), 1233zd(Z), HF, and HCl (with some minor components includingpartially fluorinated intermediates and by-products, overfluorinatedby-products, and 1233zd dimers), then enters then enters recycle columnD-1 where a stream comprising mainly unreacted1,1,1,3,3-pentachloropropane, and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene, partially fluorinated intermediates, 1233zddimers, and the majority of the HF exits the bottom of the recyclecolumn and is recycled back to the liquid phase reactor R-1 viavaporizer HX-1. A stream comprising mainly 1233zd(E), 1233zd(Z), HF, andHCl exits the top of the recycle column and enters HCl column D-2. Astream comprising mainly HCl by-product exits the top of the HCl columnand is fed to an HCl recovery system. The recovered HCl by-product canbe sold for profit. The HCl column bottoms stream consisting mainly of1233zd(E), 1233zd(Z), and HF are then fed into an HF recovery system.The HF recovery system starts with the crude 1233zd/HF stream beingvaporized in heat exchanger HX-2 and fed into HF absorption column A-1.Here a liquid stream of 50-80% H₂SO₄ contacts the gaseous 1233zd/HFstream and absorbs the majority of the HF. The stream exiting the bottomof A-1 comprises HF/H₂SO₄/H₂O and is fed to heat exchanger HX-3 where itis heated to a temperature sufficient to flash the majority of the HFalong with small amounts of H₂O and H₂SO₄. This stream is fed to HFrecovery distillation column D-2. The liquid remaining after the HF isflashed off in HX-3 consisting mainly of H₂SO₄ and H₂O (with 0-2% HF) iscooled in HX-4 and recycled back to HF absorption column A-1. The HFrecovery column, D-3, bottoms stream comprising mainly H₂SO₄ and H₂O arerecycled back to heat exchanger HX-3. Anhydrous HF is recovered from thetop of the HF recovery column, D-3, and is recycled back to the reactorR-1 via vaporizer HX-1. The stream exiting the top of HF absorptioncolumn A-1 comprising mainly 1233zd(E) and 1233zd(Z) (trace HF) is sentforward to a polishing system A-2 where the gaseous stream contacts awater or a caustic solution to remove trace HF and is subsequently driedwith a desiccant. Acid free crude product exiting absorber A-2 is sentto the first of two purification columns, D-4. A stream exiting the topof the column D-4 consists mainly of reaction by-products that haveboiling points lower than that of 1233zd(E). The stream exiting thebottom of lights column D-4 consisting mainly of 1233zd(E) and 1233zd(Z)and heavier by-products is fed to product recovery distillation columnD-5. Product grade 1233zd(E) exits the top of the column to productstorage. The product column bottoms consist mainly of 1233zd(Z) andreaction by-products with boiling points higher than that of 1233zd(E)is then fed to vaporizer HX-5 and then to isomerization reactor R-2where by-product 1233zd(Z) is converted to the desired product. Thestream leaving R-2 is then recycled to lights distillation column D-4for purification. Optionally, if any by-products in the stream enteringR-2 are unstable they may decompose and form small amounts of HF or HCl.In this case the stream exiting R-2 can be recycled and combined withthe stream entering the polishing system A-2 to remove the acid.Optionally, the stream exiting the bottom of the product recoverydistillation column, D-5 can be recycled back to liquid phase reactorR-1. In any of these options a heavies purge stream from the bottom ofthe product recovery distillation column, D-5, will be required toprevent build-up of high boiling impurities in the purification system.The heavies purge stream is collected for later use or waste disposal.

Referring to FIG. 3, shown is the synthesis of 1233zd(E) via a liquidphase reaction integrated process having a phase separation HF recoverysystem. Here, liquid phase reactor R1 is first charged with an amount ofanhydrous hydrogen fluoride that is in a stoichiometric excess of theamount needed to totally fluorinate a metal chloride catalystfluorination catalyst. E.g. when using TiCl₄ a>4:1 mole ratio of HF tocatalyst is added. This is followed by the addition of a fluorinationcatalyst alone or in combination from the group comprising TiCl₄, SnCl₄,TaCl₅, SbCl₃, FeCl₃, or AlCl₃ to prepare the catalyst. TiCl₄ is mostpreferred. The catalyst preparation is done while the reactor is at0-120° C. HCl is generated during catalyst preparation and can be ventedout of the top of the catalyst stripper column CS-1 to control thereactor pressure at or below the intended operating pressure of thereactor. Preferably the reactor is constructed from materials which areresistant to the corrosive effects of the HF and catalyst, such asHastelloy-C, Inconel, Monel, Incolloy, or fluoropolymer-lined steelvessels. Such liquid-phase fluorination reactors are well known in theart. Additional HF and 1,1,1,3,3-pentachloropropane or1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene mixture is then added until good agitation isachieved.

The reaction mixture is then heated to about 85° C. where thefluorination reaction between 1,1,1,3,3-pentachloropropane or1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene mixture and HF is initiated. Continuous1,1,1,3,3-pentachloropropane or 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene mixture andHF (in a stoichiometric excess) feeds are simultaneously fed to heaterHX-1 and then into a liquid phase reactor R-1. Optionally,1,1,1,3,3-pentachloropropane or 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene mixture isfed directly into reactor R-1 and not through heater HX-1. The operatingpressure of 60-160 psig (preferably 80-140 psig) is maintained by acontrol valve on the exiting flow from the catalyst stripper column CS-1and the reactor temperature is kept in the range of 85-120° C. primarilysupplied by steam flow into the reactor jacket. A catalyst strippercolumn CS-1 is connected to the reactor, R-1, and serves the purpose ofknocking down and returning entrained catalyst, some HF, partiallyfluorinated intermediates, and some unreacted1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene back to the reactor for further reaction.

The stream exiting the top of catalyst stripper CS-1 comprising mainly1233zd(E), 1233zd(Z), HF, and HCl (with some minor components includingpartially fluorinated intermediates and by-products, overfluorinatedby-products, and 1233zd dimers), then enters HCl column D-1. A streamcomprising mainly HCl by-product exits the top of the HCl column and isfed to an HCl recovery system. The recovered HCl by-product can be soldfor profit. The HCl column bottoms stream consisting mainly of1233zd(E), 1233zd(Z), and HF are then fed into an HF recovery system.The HF recovery system starts with the 1233zd/HF stream being fed intoheat exchanger HX-2 where it is pre-cooled to temperatures<0° C. andthen enters phase separation vessel PS-1. Here the stream temperature ismaintained or further cooled to −40-−5° C. The HF rich top layer (<10%1233zd) is recycled back to the liquid phase reactor R-1. The organicrich bottom layer containing mainly 1233zd (<4% HF) is sent to vaporizerHX-3 and then forward to a polishing system A-1 where the gaseous streamcontacts a water or a caustic solution to remove trace HF and issubsequently dried with a desiccant. Acid free crude product exitingabsorber A-1 is sent to the first of two purification columns, D-2. Astream exiting the top of the column D-2 consists mainly of reactionby-products that have boiling points lower than that of 1233zd(E). Thestream exiting the bottom of lights column D-2 consisting mainly of1233zd(E) and 1233zd(Z) and heavier by-products is fed to productrecovery distillation column D-3. Product grade 1233zd(E) exits the topof the column to product storage. The product column bottoms consistmainly of 1233zd(Z) and reaction by-products with boiling points higherthan that of 1233zd(E) is then fed to vaporizer HX-4 and then toisomerization reactor R-2 where by-product 1233zd(Z) is converted to thedesired product. The stream leaving R-2 is then recycled to lightsdistillation column D-2 for purification. Optionally, if any by-productsin the stream entering R-2 are unstable they may decompose and formsmall amounts of HF or HCl. In this case the stream exiting R-2 can berecycled and combined with the stream entering the polishing system A-1to remove the acid. Optionally, the stream exiting the bottom of theproduct recovery distillation column, D-3 can be recycled back to liquidphase reactor R-1. In any of these options a heavies purge stream fromthe bottom of the product recovery distillation column, D-3, will berequired to prevent build-up of high boiling impurities in thepurification system. The heavies purge stream is collected for later useor waste disposal.

Referring to FIG. 4, shown is the synthesis of 1233zd(E) via a liquidphase reaction integrated process having a phase separation HF recoverysystem and optional recycle column after reactor. Liquid phase reactorR1 is first charged with an amount of anhydrous hydrogen fluoride thatis in a stoichiometric excess of the amount needed to totally fluorinatea metal chloride catalyst fluorination catalyst. E.g. when using TiCl₄a>4:1 mole ratio of HF to catalyst is added. This is followed by theaddition of a fluorination catalyst alone or in combination from thegroup comprising TiCl₄, SnCl₄, TaCl₅, SbCl₃, FeCl₃, or AlCl₃ to preparethe catalyst. TiCl₄ is most preferred. The catalyst preparation is donewhile the reactor is at 0-120° C. HCl is generated during catalystpreparation and can be vented out of the top of the catalyst strippercolumn CS-1 to control the reactor pressure at or below the intendedoperating pressure of the reactor. Preferably the reactor is constructedfrom materials which are resistant to the corrosive effects of the HFand catalyst, such as Hastelloy-C, Inconel, Monel, Incolloy, orfluoropolymer-lined steel vessels. Such liquid-phase fluorinationreactors are well known in the art. Additional HF and1,1,1,3,3-pentachloropropane or 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene mixture isthen added until good agitation is achieved.

The reaction mixture is then heated to about 85° C. where thefluorination reaction between 1,1,1,3,3-pentachloropropane or1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene mixture and HF is initiated. Continuous1,1,1,3,3-pentachloropropane or 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene mixture andHF (in a stoichiometric excess) feeds are simultaneously fed to heaterHX-1 and then into a liquid phase reactor R-1. Optionally,1,1,1,3,3-pentachloropropane or 1,1,1,3,3-pentachloropropane and/or1,1,3,3-tetrachloropropene and/or 1,3,3,3-tetrachloropropene mixture isfed directly into reactor R-1 and not through heater HX-1. The operatingpressure of 60-160 psig (preferably 80-140 psig) is maintained by acontrol valve on the exiting flow from the catalyst stripper column CS-1and the reactor temperature is kept in the range of 85-120° C. primarilysupplied by steam flow into the reactor jacket. A catalyst strippercolumn CS-1 is connected to the reactor, R-1, and serves the purpose ofknocking down and returning entrained catalyst, some HF, partiallyfluorinated intermediates, and some unreacted1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene back to the reactor for further reaction.

The stream exiting the top of catalyst stripper CS-1 comprising mainly1233zd(E), 1233zd(Z), HF, and HCl (with some minor components includingpartially fluorinated intermediates and by-products, overfluorinatedby-products, and 1233zd dimers), then enters then enters recycle columnD-1 where a stream comprising mainly unreacted1,1,1,3,3-pentachloropropane and/or 1,1,3,3-tetrachloropropene and/or1,3,3,3-tetrachloropropene, partially fluorinated intermediates, 1233zddimers, and the majority of the HF exits the bottom of the recyclecolumn and is recycled back to the liquid phase reactor R-1 viavaporizer HX-1. A stream comprising mainly 1233zd(E), 1233zd(Z), HF, andHCl exits the top of the recycle column and enters HCl column D-2. Astream comprising mainly HCl by-product exits the top of the HCl columnand is fed to an HCl recovery system. The recovered HCl by-product canbe sold for profit. The HCl column bottoms stream consisting mainly of1233zd(E), 1233zd(Z), and HF are then fed into an HF recovery system.The HF recovery system starts with the 1233zd/HF stream being fed intoheat exchanger HX-2 where it is pre-cooled to temperatures<0° C. andthen enters phase separation vessel PS-1. Here the stream temperature ismaintained or further cooled to −40-0° C. The HF rich top layer (<10%1233zd) is recycled back to the liquid phase reactor R-1. The organicrich bottom layer containing mainly 1233zd (<4% HF) is sent to vaporizerHX-3 and then forward to a polishing system A-1 where the gaseous streamcontacts a water or a caustic solution to remove trace HF and issubsequently dried with a desiccant. Acid free crude product exitingabsorber A-1 is sent to the first of two purification columns, D-3. Astream exiting the top of the column D-3 consists mainly of reactionby-products that have boiling points lower than that of 1233zd(E). Thestream exiting the bottom of lights column D-3 consisting mainly of1233zd(E) and 1233zd(Z) and heavier by-products is fed to productrecovery distillation column D-4. Product grade 1233zd(E) exits the topof the column to product storage. The product column bottoms consistmainly of 1233zd(Z) and reaction by-products with boiling points higherthan that of 1233zd(E) is then fed to vaporizer HX-4 and then toisomerization reactor R-2 where by-product 1233zd(Z) is converted to thedesired product. The stream leaving R-2 is then recycled to lightsdistillation column D-3 for purification. Optionally, if any by-productsin the stream entering R-2 are unstable they may decompose and formsmall amounts of HF or HCl. In this case the stream exiting R-2 can berecycled and combined with the stream entering the polishing system A-1to remove the acid. Optionally, the stream exiting the bottom of theproduct recovery distillation column, D-4 can be recycled back to liquidphase reactor R-1. In any of these options a heavies purge stream fromthe bottom of the product recovery distillation column, D-4, will berequired to prevent build-up of high boiling impurities in thepurification system. The heavies purge stream is collected for later useor waste disposal.

Vapor Phase Reaction

The reaction chemistry for the vapor phase process involves asingle-step reaction of 1,1,1,3,3-pentachloropropane (HCC-240fa) withanhydrous HF in a vapor phase reactor to produce a mixture ofHCFO-1233zd (1-chloro-3,3,3-trifluoropropene) and HCl. Preferably,though not exclusively, the reaction is maintained under conditions(e.g. temperature, pressure, residence time, etc.) to increase therelative ratio of (E) to (Z) isomers of HCFO-1233zd. The instantreaction also minimizes the undesirable reaction of HF with theresulting HCFO-1233zd(E), which produces HCF-244fa and/or HCF-245fa,either of which proceeds to the formation of HFO-1234ze (as illustratedbelow).

Accordingly, in one embodiment, the desired reaction of the instantinvention is illustrated as follows:

In a second embodiment, the undesired reactions of the instant inventionare illustrated as follows:

In certain embodiments, the manufacturing process comprises five majorunit operations: (1) continuous mode fluorination reaction using HF, (2)separation and purification of at least byproduct HCl, (3) separation ofexcess HF back to (1), (4) production/purification of final product,HCFO-1233zd(E) with optional recycle of high-boilers back to (1), and(5) isomerization of by-product HCFO-1233zd(Z) to 1233zd(E). Thismaximizes process yield. The relative positions of these operations areshown in FIGS. 5-8, and are discussed in greater detail below.

Unit Operation One: Catalytic fluorination of1,1,1,3,3-pentachloropropane

The fluorination reaction described herein uses a vapor phase catalystof sufficient strength and reaction conditions to achieve the desiredreaction product, i.e. HCFO-1233zd(E)

The fluorination catalyst may include any vapor phase fluorinationcatalyst that is known in the art. Suitable catalysts include, but arenot limited to, chromium, aluminum, cobalt, manganese, nickel and ironoxides, hydroxides, halides, oxyhalides, inorganic salts thereof andtheir mixtures, which may be supported or in bulk. The vapor phasefluorination catalyst may also include combinations of known catalysts,such as, but not limited to, Cr₂O₃, Cr₂O₃/Al₂O₃, Cr₂O₃/AlF₃,Cr₂O₃/carbon, CoCl₂/Cr₂O₃/Al₂O₃, NiCl₂/Cr₂O₃/Al₂O₃, CoCl₂/AlF₃,NiCl₂/AlF₃ and mixtures thereof. Additional fluorination catalysts thatcan be used include, but are also not limited to, FeCl₃/C, SnCl₄/C,TaCl₅/C, SbCl₃/C, AlCl₃/C, and AlF₃/C. Supports for the metal halideslisted can also be alumina or fluorinated alumina. Any of the foregoingcatalysts, as well as other catalysts not listed, may be partially ortotally fluorinated by anhydrous HF.

In certain embodiments, the fluorination catalyst may include chromium(III) oxides, such as crystalline chromium oxide or amorphous chromiumoxide. While not limited thereto, amorphous chromium oxide (Cr₂O₃) ismost preferred vapor phase catalyst. It is a commercially availablematerial in a variety of particle sizes and may be selected to enhancetheir effectiveness. In certain embodiments, it is provided having apurity of at least 98%. The fluorination catalyst is provided in anyamount sufficient to drive the reaction, but also may be presented inexcess.

This reaction may be conducted in a vapor phase reactor known in theart. Preferably, though not exclusively, the reactor is constructed frommaterials that are resistant to the corrosive effects of the HF andcatalyst, such as Hastelloy-C, Inconel, Monel, Incalloy, orfluoropolymer-lined steel vessels. Such vapor-phase fluorinationreactors are well known in the art.

Proper temperature and pressure control of the reaction are desirablefor optimum conversion. In certain non-limiting embodiments, thereaction temperature should be about 200 to about 450° C. and thereaction pressure should be about 0 to about 160 psig pressure. Infurther embodiments, the reaction temperature should be between about250 and 400° C. and the reaction pressure should be about 0 to about 140psig pressure. In even further embodiments, the reaction temperatureshould be between about 275 and 375° C. and the reaction pressure shouldbe about 2 to about 130 psig pressure. It has been surprisinglydiscovered that maintaining the reaction under these operatingconditions, produces an unexpected shift in the reaction mechanism whichproduces a high ratio of HCFO-1233zd(E) compared to HCFO-1233zd(Z).

Unit Operation Two: Removal of HCl

The HCl containing reaction products stream continuously exits thereactor and flows into an attached HCl distillation column. The HClmaterial exits the top of the column and can then be further purified.High purity HCl can be isolated and absorbed in de-ionized water asconcentrated HCl for other uses.

Unit Operation Three: Separation and Recycle of Excess HF

The bottom stream from the HCl removal column that contains crudeproduct mixture of HCFO-1233zd(E) and HF (in some embodiments about 30wt % to about 60 wt %) is fed to a sulfuric acid extractor or phaseseparator for removal of HF from this mixture. HF is either dissolved inthe sulfuric acid or phase separated from the organic mixture. Forembodiments utilizing a sulfuric acid adsorption system, the HF may bedesorbed from the sulfuric acid/HF mixture by stripping distillation andrecycled back to the reactor. For embodiments utilizing a phaseseparator, HF is phase-separated and recycled back to the reactor. Theorganic mixture either from the overhead of the sulfuric acid extractoror from the bottom layer of the phase separator may employ treatment(e.g. scrubbing or adsorption) to remove traces of HF before it is fedto the next unit operation (purification of the final product).

Unit Operation Four: Purification of the Final Product

Purification of the final product may include the use of one or moredistillation columns.

In certain embodiments, purification includes two or more continuouslyoperating distillation columns. The first column is used to remove thelighter components from the mixture, such as HFO-1234ze. The secondcolumn may be used to removed heavier components, e.g. HCFO-1233zd(Z)and purify the final product, HCFO-1233zd(E). Such heavier components,if desirable, may be fed to an isomerization reactor, collected forfurther use, or optionally recycled back to the reactor. In certainembodiments, it is desirable to have a purge of heavy by-products fromthis stream.

Unit Operation Five: Isomerization of by-Product HCFO-1233zd(Z) toHCFO-1233zd(E)

To maximize the HCFO-1233zd(E) yield in this process, the by-productHCFO-1233zd(Z) formed in the reaction and exiting the second column isfed as a vapor to a reactor that contains isomerization catalyst. Theby-product is then converted to the desired product. The isomerizationreactor exit stream is then recycled for purification, using theforegoing method. The isomerization reactor can be either thefluorination reactor discussed above or a separate isomerizationreactor.

Catalysts may include, but are not limited to, any of the vapor phasecatalysts discussed herein. In certain non-limiting embodiments, thecatalyst is a fluorinated chromium oxide.

In certain preferred embodiments, this step involves controlling thetemperature of a heated surface to greater than 50° C.-375° C. Theheated surface is contacted with the stream containing theHCFO-1233zd(Z) by-product. The feed stream is contacted with the heatedsurface for a period of time sufficient to convert at least a portion ofthe 1233zd(Z) to 1233zd(E) to produce a product stream rich in1233zd(E).

In some embodiments, the heated surface includes the inside of a reactorvessel. In addition, or in the alternative, the heated surface mayinclude an outer surface of a packing material, for example a packingmaterial that is packed in a reaction vessel. In some embodiments, thereactor vessel is a batch-wise reactor vessel that can be charged withthe feed stream. In some such embodiments, the feed stream may be sealedin the batch-wise reactor, and, after sufficient time passes toisomerize the desired amount of HCFO-1233zd(Z), the reactor vessel maybe opened to remove the product stream. In other embodiments, thereactor vessel is a continuous-type reactor vessel, for example areactor vessel with a first opening and a second opening and a fluidpathway between the first and second openings. The feed stream is fedinto the reactor vessel through the first opening and passes through thereactor vessel at a rate sufficient to isomerize the desired amount ofHCFO-1233zd(Z). The resulting product stream exits the second opening.In one example, the reactor vessel is an elongate reactor vessel (e.g.,a MONEL™ tube) with the first opening at a first end and the secondopening at a second end.

In some embodiments, the reactor vessel may be partially or entirelypacked with packing material, for example with a stainless steelpacking. In some embodiments, the relatively large surface area of thepacking material may facilitate the conversion reaction from the (Z) tothe (E) isomer. Support structures that support the packing material mayalso be disposed in or on the reactor vessel. For example, the packingmaterial may be supported by a mesh or other structure that is disposedunder, around, and/or within the packing material. The support structuremay comprise the same material as the packing material (e.g., stainlesssteel), nickel, or any other suitable material.

The packing materials may also comprise one or more catalyst materials.Examples of suitable catalysts for the isomerization of HCFO-1233zd aremetal oxides, halogenated metal oxides, Lewis acid metal halides,zero-valent metals, as well as combinations of these catalysts. Specificexamples of suitable catalysts are AlF₃, Cr₂O₃, fluorinated Cr₂O₃,zirconium oxide and halogenated versions thereof, or an aluminum oxideand halogenated versions thereof. In addition, the catalysts may beactivated prior to use. Examples of activation procedures for severalsuitable catalysts may be found in U.S. Publication No. 2008-0103342,which is hereby incorporated by reference in its entirety.

Referring to FIG. 5 the synthesis of HCFO-1233zd(E) via a vapor phasereaction integrated process having a sulfuric acid HF recovery isillustrated. The vapor phase reactor 10 is first loaded with afluorination catalyst, as discussed above. Catalysts can be supported orin bulk, with preferred, though non-limiting catalysts, beingfluorinated chromium oxide.

HCC-240fa and HF are simultaneously fed to a vaporizer 12 and then intothe vapor phase reactor 10. The reaction temperature may be about 200 toabout 450° C. and at about 0 to about 160 psig pressure. The mole ratioof HF to HCC-240fa may be 3:1, preferably between 3:1 and 20:1, morepreferably between 4:1 and 12:1, and most preferably between 5:1 and11:1.

Reactor effluent including partially fluorinated intermediates andby-products, overfluorinated by-products, HF, HCFO-1233zd(E+Z) and HCl,then enters HCl column 14 through line 16. A stream of mainly HClby-product exits the top portion of the HCl column 14 and is fed to anHCl recovery system through line 18. The recovered HCl by-product can beused for other purposes, as discussed herein.

The HCl column bottoms consisting mainly of partially fluorinatedintermediates and by-products, overfluorinated by-products, HF andHCFO-1233zd(E+Z) are then fed via line 20 into an HF recovery system.The HF recovery system starts with the HCFO-1233zd/HF stream beingvaporized in heat exchanger 22 and is then fed into HF absorption column24 via line 26. Here, a liquid stream of 50-80% H₂SO₄ contacts thegaseous HCFO-1233zd/HF stream and absorbs the majority of the HF. Thestream 28 exiting the bottom of column 24 includes HF, H₂SO₄ and H₂O andis fed to heat exchanger 30 where it is heated to a temperaturesufficient to flash the majority of the HF along with small amounts ofH₂O and H₂SO₄. This stream is fed via line 32 to HF recoverydistillation column 34. The liquid remaining after the HF is flashed offin heat exchanger 30 including mainly H₂SO₄ and H₂O (with 0-2% HF) iscooled in heat exchanger 36 via line 38 and recycled via line 40 back toHF absorption column 24. The HF recovery column 34 bottoms streamincluding mainly H₂SO₄ and H₂O is recycled via line 42 back to heatexchanger 30.

Anhydrous HF is recovered from the top of the HF recovery column 34 andis recycled back to the reactor 10 via line 44 to vaporizer 12.

The stream exiting the top portion of HF absorption column 24 includingmainly HCFO-1233zd (E+Z) (trace HF) is sent forward via line 48 to apolishing system 46 where the gaseous stream contacts water or a causticsolution to remove trace HF and is subsequently dried with a desiccant.Acid free crude product exiting absorber 46 is sent via line 50 to thefirst of two purification columns 52. A stream 54 exiting the topportion of the column 52 includes mainly of reaction by-products thathave boiling points lower than that of HCFO-1233zd(E). The stream 56exiting the bottom portion of lights column 52 consisting mainly ofHCFO-1233zd(E+Z) and heavier by-products is fed to product recoverydistillation column 58. Product grade HCFO-1233zd(E) exits the topportion of the column to product storage via line 60. The product columnbottoms 62 includes mainly HCFO-1233zd(Z).

Reaction by-products with boiling points higher than that ofHCFO-1233zd(E) are then fed to a vaporizer (not illustrated) and then toisomerization reactor (not illustrated) where by-product HCFO-1233zd(Z)is converted to the desired product. The stream leaving is then recycledto lights distillation column 52 for purification. Optionally, if anyby-products in the stream entering are unstable, they may decompose andform small amounts of HF or HCl. In this case, the stream exiting can berecycled and combined with the stream entering the polishing system toremove the acid. Optionally, the stream exiting the bottom of theproduct recovery distillation column 58 can be recycled back to vaporphase reactor 10, where isomerization of Z to E isomer of HCFO-1233zdtakes place. In any of these possibilities, a heavies purge stream fromthe bottom of the product recovery distillation column 58 preventsbuild-up of high boiling impurities in the purification system. Theheavies purge stream is collected for later use or waste disposal. Afterdeactivation of the catalyst in reactor 10 it can be regenerated in situby heating to 300-400° C. and passing an oxidizing agent such as O₂ orCl₂ over the catalyst for a selected period of time.

Referring to FIG. 6, shown is the synthesis of HCFO-1233zd(E) via avapor phase reaction integrated process with sulfuric acid, HF recoveryand optional recycle column after reactor. Specifically, a vapor phasereactor 100 is first loaded with a fluorination catalyst, as discussedabove. Catalysts can be supported or in bulk, with preferred, thoughnon-limiting catalysts, being fluorinated chromium oxide.

HCC-240fa and HF are simultaneously fed to a vaporizer 112 and then intoa vapor phase reactor 100 via line 114. The reaction temperature may beabout 200 to about 450° C. and at about 0 to about 160 psig pressure.The mole ratio of HF to HCC-240fa may be 3:1, preferably between 3:1 and20:1, more preferably between 4:1 and 12:1, and most preferably between5:1 and 11:1. The preferred catalyst in 100 is fluorinated chrome oxide.

The reactor effluent including partially fluorinated intermediates andby-products, overfluorinated by-products, HF, HCFO-1233zd(E+Z) and HCl,then enters recycle column 116 via line 118 where a stream 120 includingmainly unreacted HCC-240fa, partially fluorinated intermediates, and themajority of the HF exits the bottom portion of the recycle column 116and is recycled back to the vapor phase reactor 100 via vaporizer 112. Astream including mainly HCFO-1233zd(E), HF, and HCl exits the topportion of the recycle column and enters HCl column 122 via line 124. Astream 126 including mainly HCl by-product exits the top of the HClcolumn 122 and is fed to an HCl recovery system. The recovered HClby-product can be used for other purposes, as discussed herein. The HClcolumn bottoms including mainly partially fluorinated by-products,overfluorinated by-products, HF and HCFO-1233zd(E+Z) are then fed intoan HF recovery system via line 128.

The HF recovery system starts with the crude HCFO-1233zd/HF stream beingvaporized in heat exchanger 130 and fed into HF absorption column 132.Here, a liquid stream of 50-80% H₂SO₄ contacts the gaseousHCFO-1233zd/HF stream and absorbs the majority of the HF. The streamexiting the bottom of column 132 includes HF/H₂SO₄/H₂O and is fed toheat exchanger 136 via line 138 where it is heated to a temperaturesufficient to flash the majority of the HF along with small amounts ofH₂O and H₂SO₄. This stream is fed to HF recovery distillation column140. The liquid remaining after the HF is flashed off in heaterexchanger 136 includes mainly H₂SO₄ and H₂O (with 0-2% HF) and is cooledin heat exchanger 142 via line 144 and recycled back to HF absorptioncolumn 132 via line 146.

The HF recovery column 140 bottoms stream including mainly H₂SO₄ and H₂Ois recycled back to heat exchanger 136 via line 148. Anhydrous HF isrecovered from the top portion of the HF recovery column 140, and isrecycled back to the reactor 100 via line 150 and heat exchanger 112.The stream exiting the top of HF absorption column 132 including mainlyHCFO-1233zd (E+Z) (trace HF) is sent forward via line 154 to a polishingsystem 152 where the gaseous stream contacts a water or a causticsolution to remove trace HF and is subsequently dried with a desiccant.Acid free crude product exiting absorber 152 is sent to the first of twopurification columns 156 (and subsequently 162) via line 158. A stream166 exiting the top portion of the column 156 includes mainly reactionby-products that have boiling points lower than that of HCFO-1233zd(E).The stream exiting the bottom portion of lights column 156 includesmainly HCFO-1233zd(E+Z) and heavier by-products is fed via line 160 toproduct recovery distillation column 162. Product grade HCFO-1233zd(E)exits the top portion of column 162 via line 164 to product storage. Theproduct column bottoms 168 include mainly HCFO-1233zd(Z) and reactionby-products with boiling points higher than that of HCFO-1233zd(E) arethen fed to vaporizer (not illustrated) and then to isomerizationreactor (not illustrated) where by-product HCFO-1233zd(Z) is convertedto the desired product. The stream leaving is then recycled to lightsdistillation column for purification. Optionally, if any by-products inthe stream entering are unstable, they may decompose and form smallamounts of HF or HCl. In this case, the stream exiting can be recycledand combined with the stream entering the polishing system to remove theacid. Optionally, the stream exiting the bottom portion of the productrecovery distillation column can be recycled back to the vapor phasereactor where isomerization of Z to E isomer of HCFO-1233zd takes place.In any of these options, a heavies purge stream from the bottom of theproduct recovery distillation column prevents build-up of high boilingimpurities in the purification system. The heavies purge stream iscollected for later use or waste disposal. After deactivation of thecatalyst in reactor 100, it can be regenerated in situ by heating toabout 300 to about 400° C. and passing an oxidizing agent such as O₂ orCl₂ over it for a selected period of time.

Referring to FIG. 7, shown is the synthesis of HCFO-1233zd(E) via avapor phase reaction integrated process having a phase separation HFrecovery system. Here, a vapor phase reactor 210 is first loaded with afluorination catalyst, as discussed above. Catalysts can be supported orin bulk, with preferred, though non-limiting catalysts, beingfluorinated chromium oxide.

HCC-240fa and HF are simultaneously fed to a vaporizer 200 and then intoa vapor phase reactor 210 via line 212. The reaction temperature may beabout 200 to about 450° C. and at about 0 to about 160 psig pressure.The mole ratio of HF to HCC-240fa is 3:1, preferably between 3:1 and20:1, more preferably between 4:1 and 12:1, and most preferably between5:1 and 11:1. The preferred catalyst in reactor 210 is fluorinatedchrome oxide. The reactor effluent includes partially fluorinatedintermediates and by-products, overfluorinated by-products,overfluorinated by-products, HF, HCFO-1233zd(E+Z) and HCl, then entersHCl column 214 via line 216. A stream 218 including mainly HClby-product exits the top portion of the HCl column 214 and is fed to anHCl recovery system (not shown) via line 218. The recovered HClby-product can be used for other purposes, as discussed herein. The HClcolumn bottoms consisting mainly of partially fluorinated intermediatesand by-products, overfluorinated by-products, HF and HCFO-1233Z(E+Z) arethen fed via line 220 into an HF recovery system. The HF recovery systemstarts with the HCFO-1233zd/HF stream being fed into heat exchanger 221where it is pre-cooled to temperatures≦0° C. and then enters phaseseparation vessel 222 via line 224. The stream temperature is maintainedor further cooled to about −40 to about 0° C. The HF rich top layer(<10% 1233zd) is recycled via line 226 back to the vapor phase reactor210. The organic rich bottom layer containing mainly HCFO-1233zd (<4%HF) is sent via line 228 to vaporizer 230 and the forward to a polishingsystem 232 where the gaseous stream contacts a water or a causticsolution to remove trace HF and is subsequently dried with a desiccant.Acid free crude product exiting absorber is sent via line 234 to thefirst of two purification columns 236. A stream 238 exiting the topportion of the column 236 includes mainly reaction by-products that haveboiling points lower than that of HCFO-1233zd(E). The stream 240 exitingthe bottom of lights column 236 including mainly HCFO-1233zd(E+Z) andheavier by-products is fed to product recovery distillation column 242.Product grade HCFO-1233zd(E) exits the top of the column 242 to productstorage via line 244. The product column bottoms 246 include mainlyHCFO-1233zd(Z) and reaction by-products with boiling points higher thanthat of HCFO-1233zd(E) is then fed to the vaporizer and then to theisomerization reactor where by-product HCFO-1233zd(Z) is converted tothe desired product. The stream leaving is then recycled to the lightsdistillation column for purification. Optionally, if any by-products inthe stream entering are unstable, they may decompose and form smallamounts of HF or HCl. In this case, the stream exiting can be recycledand combined with the stream entering the polishing system to remove theacid. Optionally, the stream exiting the bottom portion of the productrecovery distillation column 242 can be recycled back to the vapor phasereactor where isomerization of Z to E isomer of HCFO-1233zd takes place.In any of these options, a heavies purge stream from the bottom portionof the product recovery distillation column prevents build-up of highboiling impurities in the purification system. The heavies purge streamis collected for later use or waste disposal. After deactivation of thecatalyst in reactor 210 it can be regenerated in situ by heating toabout 300 to about 400° C. and passing an oxidizing agent such as O₂ orCl₂ over it for a prescribed period of time.

Referring to FIG. 8, shown is the synthesis of HCFO-1233zd(E) via avapor phase reaction integrated process having a phase separation HFrecovery system and optional recycle column after reactor. Morespecifically, a vapor phase reactor 300 is first loaded with afluorination catalyst, as discussed above. Catalysts can be supported orin bulk, with preferred, though non-limiting catalysts, beingfluorinated chromium oxide.

HCC-240fa and HF are simultaneously fed to a vaporizer 310 and then intovapor phase reactor 300 via line 312. The reaction temperature may beabout 200 to about 450° C. and at about 0 to about 160 psig pressure.The mole ratio of HF to HCC-240fa is 3:1, preferably between 3:1 and20:1, more preferably between 4:1 and 12:1, and most preferably between5:1 and 11:1. The preferred catalyst in reactor 300 is fluorinatedchrome oxide. The reactor effluent including partially fluorinatedintermediates and by-products, overfluorinated by-products, HF,HCFO-1233zd(E+Z) and HCl, then enters recycle column 314 via line 316where a stream including mainly unreacted HCC-240fa, partiallyfluorinated intermediates, and the majority of the HF exits the bottomportion of recycle column 314 and is recycled back to the vapor phasereactor 300 via line 318 and vaporizer 310. A stream including mainlyHCFO-1233zd(E), HF and HCl exits the top portion of the recycle column314 via line 320 and enters HCl column 322. A stream 324 includingmainly HCl by-product exits the top portion of HCl column 322 and is fedto an HCl recovery system (not shown). The recovered HCl by-product canbe used for other purposes, as discussed herein. The HCl column bottomsincluding mainly partially fluorinated by products, overfluorinatedby-products, HF and HCFO-1233zd(E+Z) are then fed via line 326 into anHF recovery system.

The HF recovery system starts with the HCFO-1233zd/HF stream being fedinto heat exchanger 328 where it is pre-cooled to temperatures<0° C. andthen enters phase separation vessel 330 via line 332. The streamtemperature may be maintained or further cooled to about −40 to about 0°C. The HF rich top layer (<10% 1233zd) may be recycled back to the vaporphase reactor 300 via line 334. The organic rich bottom layer containingmainly HCFO-1233zd (<4% HF) is sent via line 336 to vaporizer 338 andthen forwarded to a polishing system 340 where the gaseous streamcontacts a water or a caustic solution to remove trace HF and issubsequently dried with a desiccant.

Acid free crude product exiting absorber 340 is sent to the first of twopurification columns, 342 via line 344. A stream 346 exiting the top ofthe column 342 includes mainly reaction by-products that have boilingpoints lower than that of HCFO-1233zd(E). The stream 348 exiting thebottom portion of lights column 342 including mainly HCFO-1233zd(E+Z)and heavier by-products is fed to product recovery distillation column350. Product grade HCFO-1233zd(E) exits the top of the column to productstorage via line 352. The product column bottoms 354 include mainlyHCFO-1233zd(Z) and reaction by-products with boiling points higher thanthat of HCFO-1233zd(E) are then fed to the vaporizer (not illustrated)and then to the isomerization reactor (not illustrated) where by-productHCFO-1233zd(Z) is converted to the desired product. The stream leavingis then recycled to the lights distillation column for purification.Optionally, if by-products in the stream entering are unstable, they maydecompose and form small amounts of HF or HCl. In this case, the streamexiting can be recycled and combined with the stream entering thepolishing system to remove the acid. Optionally, the stream exiting thebottom portion of the product recovery distillation column can berecycled back to the vapor phase reactor where isomerization of Z to Eisomer of HCFO-1233zd takes place. In any of these options, a heaviespurge stream from the bottom portion of the product recoverydistillation column prevents build-up of high boiling impurities in thepurification system. The heavies purge stream is collected for other useor waste disposal. After deactivation of the catalyst in reactor 300 itcan be regenerated in situ by heating to about 300 to about 400° C. andpassing an oxidizing agent such as O₂ or Cl₂ over it for a selectedperiod of time.

Specific embodiments of the present invention will now be described inthe following Examples. The Examples are illustrative only, and are notintended to limit the remainder of the disclosure in any way.

EXAMPLES Example 1

This example (called Run #3) illustrates the semi-batch, liquid phase,reaction where HF was continuously fed into a charge of Titaniumtetrachloride catalyst and 1,1,1,3,3-pentachloropropane (HCC-240fa).

A clean, empty 10-gallon jacketed, agitated reactor of Hastelloy Cconstruction was prepared. This reactor was connected to a 2″ vertical,PTFE-lined pipe containing packing material (stripper), which was inturn connected to an overhead heat exchanger. The heat exchanger wassupplied with −40° C. brine circulation on the shell side. Vaporsexiting this stripper were processed through a scrubber, in whichtemperature-controlled dilute potassium hydroxide aqueous solution wascirculated. Vapors exiting this stripper were collected in a weighed,chilled (−40° C.) cylinder referred to as the product collectioncylinder, followed by a smaller cylinder in series chilled in a dry icebath.

For Run #3, 14 lbs. of anhydrous HF was fed to assure catalystfluorination. Next, 1.5 lbs. of TiCl₄ was added as a catalyst. HCl wasimmediately generated as observed by the build-up of pressure in thereactor. After the pressure was reduced by venting most of the HCl fromthe system, 50 lbs. of HCC-240fa was added. The reactor was heated. Atabout 85° C. HCl started to be generated indicating that thefluorination reaction was initiated. The system pressure was controlledat about 120 psig. Additional HF was then fed continuously and productwas collected in the product collection cylinder until the HCC-240fa wasconsumed.

The GC analysis of the crude material collected during the run was asfollows:

TABLE 1 Wt. Percent Compound 86.4% 1233zd(E) 5.5% G-244fa 3.1% 1234ze(E)1.5% 1233zd(Z) 1.1% 1234ze(Z) 1.1% dimer 0.2% trifluoropropyne

Example 2

Following Run #3 from Example 1, the reactor was drained, and a freshcharge of catalyst was made, and Run #4 was performed in a similarmanner to Run #3.

The catalyst charge for Run #4 was 753 grams of TiCl₄. The operatingscheme for this run continued the same as Run #3—an initial batch chargeof 14.3 lbs of HF was added before the catalyst. Then HCC-240fa (51.9lbs) was added on top of the catalyst after HCl from the fluorination ofthe catalyst was complete. After the reaction temperature was achieved,continuous HF feed was started and maintained until either the analysisof the product quality showed a dramatic drop, or the collection ofproduct weight subsided—at these times, a fresh charge of HCC-240fa wasmade (typically 25-30 lbs.) and continuous HF feed was resumed.

Catalyst productivity for Run #4 was approximately 0.4 pph/lb ofcatalyst. About 244 lbs of crude 1233zd(E) was collected during the run.

The run was continued for over 900 hours without loss of catalystactivity.

Example 3

This example illustrates a continuous liquid phase fluorination reactionwhere HF and an organic feed mixture of 1,1,1,3,3-pentachloropropane,1,1,3,3-tetrachloropropene, and 1,3,3,3-tetrachloropropene iscontinuously fed into a charge of Titanium tetrachloride catalyst.

The same 10-gallon jacketed reactor system from Examples 1 and 2 isused.

30 lbs. of anhydrous HF is charged to the reactor. This amount is inexcess of that needed to fluorinate the TiCl4 catalyst. The agitator isstarted at 250 RMPs. Next, 1.5 lbs. of TiCl₄ is added as a catalyst. HClis immediately generated as observed by the build-up of pressure in thereactor. After the pressure is reduced by venting most of the HCl fromthe system, 20 lbs. of an organic feed mixture consisting of 70 wt %1,1,1,3,3-pentachloropropane, 27 wt % 1,1,3,3-tetrachloropropene, and 3wt % 1,3,3,3-tetrachloropropene is added. The reactor is then heated byadding steam to the jacket. At about 85° C. HCl starts to be generatedindicating that the fluorination reaction has been initiated. The systempressure is then controlled at about 120 psig. Additional HF and organicfeed mixture is then fed continuously at a mole ratio of 9:1 HF toorganic. A stream consisting mainly of HCl by-product, 1233zd, andexcess HF is vented off the top of the catstripper, scrubbed to removeacid, and organic collected in the product collection cylinder.

The GC analysis of the crude material collected during the run was asfollows:

TABLE 2 Wt. Percent Compound  94% 1233zd(E) 3.2% G-244fa 1.4% 1234ze(E)0.5% 1233zd(Z) 0.5% others

Example 4

This example illustrates the recovery of anhydrous HF from a mixture ofHF and HCFO-1233zd according to certain preferred embodiments of thepresent invention.

A mixture consisting of about 70 wt. % HCFO-1233zd(E) crude and about 30wt. % HF is vaporized and fed to the bottom of a packed column at a feedrate of about 2.9 lbs per hour for about 4 hours. A stream of about 80wt. % sulfuric acid (80/20 H₂SO₄/H₂O) with about 2% HF dissolved thereinis fed continuously to the top of the same packed column at a feed rateof about 5.6 lbs per hour during the same time frame. A gaseous streamexiting the top of the column comprises HCFO-1233zd(E) crude with lessthan 1.0 wt. % HF therein. The concentration of HF in the sulfuric acidin the column bottoms increases from 2.0 wt. % to about 15 wt. %.

The column bottoms containing sulfuric acid and about 15 wt. % HF iscollected and charged into a 2 gallon teflon vessel. The mixture isheated to about 140° C. to vaporize and flash off HF product, which iscollected. The collected HF product contains about 6000 ppm water and500 ppm sulfur. The sulfuric acid contains about 500 ppm of TOC (totalorganic carbon).

The HF collected from flash distillation is distilled in a fractionationdistillation column and anhydrous HF is recovered. The recoveredanhydrous HF contains less than 50 ppm of sulfur impurities and lessthan 100 ppm water.

Example 5

This example demonstrates the purification of the acid free 1233zd(E)crude product.

About 92 lbs of acid free 1233zd crude material produced in Example 2was charged to a batch distillation column. The crude material containedabout 94 GC area % 1233zd(E) and 6 GC area % impurities. Thedistillation column consisted of a 10 gallon reboiler, 2 inch ID by 10feet propack column, and a shell and tube condenser. The column hadabout 30 theoretical plates. The distillation column was equipped withtemperature, pressure, and differential pressure transmitters. About 7lbs of a lights cut was recovered which consisted of mainly 1234ze(Z+E),trifluoropropyne, 245fa, and 1233zd(E). 82 lbs of 99.8+GC area %1233zd(E) were collected. The reboiler residue amounting to about 3 lbswas mainly 244fa, 1233zd(Z), 1233zd dimmer, and 1233zd(E). The recoveryof 99.8+GC area % pure 1233zd(E) was 94.8%.

Example 6

This example demonstrates the purification of the acid free 1233zd(E)crude product. About 92 lbs of acid free 1233zd crude material producedin Example 2 was charged to a batch distillation column. The crudematerial contained about 94 GC area % 1233zd(E) and 6 GC area %impurities. The distillation column consisted of a 10 gallon reboiler, 2inch ID by 10 feet propack column, and a shell and tube condenser. Thecolumn had about 30 theoretical plates. The distillation column wasequipped with temperature, pressure, and differential pressuretransmitters. About 7 lbs of a lights cut was recovered which consistedof mainly 1234ze(Z+E), trifluoropropyne, 245fa, and 1233zd(E). 82 lbs of99.8+GC area % 1233zd(E) were collected. The reboiler residue amountingto about 3 lbs was mainly 244fa, 1233zd(Z), 1233zd dimmer, and1233zd(E). The recovery of 99.8+GC area % pure 1233zd(E) was 94.8%.

Example 7

This example demonstrates the use of the optional recycle column.

A representative 1233zd(E) liquid phase reactor effluent mixture asdetermined in Example 2 is charged into a batch distillation column. Thedistillation column consists of a 10 gallon reboiler, 2 inch ID by 10feet propack column, and a shell and tube condenser with −40° C. coolantflow capability. The column has about 30 theoretical plates. Thedistillation column is equipped with temperature, pressure, anddifferential pressure transmitters. The distillation column feed mixtureis about 30 wt % HF, 37 wt % HCl and 33% 1233zd(E) crude. Thedistillation is run at a pressure of about 100 psig and a differentialpressure (delta P) of 15-20 inches of water. Both the distillate andreboiler are sampled periodically and analyzed for organic, HF, and HClusing gas and ion Chromatography. Initially, HCl, organic, and HF areobserved in both samples. As more material is removed as distillate theconcentration of the reboiler changes. First, the concentration of HCldecreases until it is undetectable. The distillation is allowed toproceed until the concentration of organic in the reboiler sampledecreases to only trace amounts as analyzed using gas chromatography. Atthe conclusion of the distillation the material remaining in thereboiler is essentially pure HF. The recovered HF (reboiler bottoms) isthen used to demonstrate recycle of recovered HF back to the liquidphase fluorination reactor and works satisfactorily.

Example 8

This example demonstrates the HF recovery by phase separation.

It is visually observed using a Teflon cylinder that HF and 1233zd(E)form a heterogeneous mixture. The separation of 1233zd and HF layers istested in the temperature range from +10° C. to −30° C. Thephase-separation of a mixture containing 1233zd(E) and HF is performedin the temperature range of −30° C. to +10° C. A 500 ml SS samplecylinder is used for the study. The temperature of the cylinder iscontrolled with ethanol circulating through the coil wrapped around thecylinder. A thermocouple is attached to the outside wall of the cylinder(between cooling coil and the cylinder wall) and positioned in themiddle of the cylinder to measure the temperature. The cylinder is alsoequipped with sampling valves at the bottom and the top of the cylinder.To the cylinder is charged 100 g of anhydrous HF and 250 g of a1233zd(E). The weight ratio HF:1233zd(E) is 28.6:71.4. The cylinder ispadded with nitrogen to 15 psig at −30° C. to allow sampling. Samplesare taken from the bottom of the cylinder into Tedlar gas sample bagsthat contains 5 grams of distilled water for the purpose of absorbingHF. The first sample is taken two hours after the cylinder reaches thedesired temperature. HF concentration is determined by titration with0.1 N KOH of the aqueous phase of the sample bags. HF concentration insamples taken after 2 hours at given temperature is presented in Table3.

HF concentration in the HF layer is analyzed after the organic layer wasremoved from the system. KOH titration showed that concentration of HFin the acid layer was about 70±5%.

TABLE 3 HF concentration in the samples of the bottom (organic) phasetaken after equilibrating the contents of the phase-separator for 2hours at given temperature HF concentration in bottom Temperature (° C.)(organic phase (wt %) −30 1.00 −20 1.25 −10 2.75 0 3.25 10 4.00

Example 9

This example demonstrates the isomerization of 1233zd(Z) into desiredproduct 1233zd(E).

Conversion of 1233zd(Z) into 1233zd(E) was performed using a MONEL™reactor (ID 2 inch, length 32 inch) equipped with a MONEL™ preheater (ID1 inch, length 32 inch) which was filled with Nickel mesh to enhanceheat transfer. The reactor was filled with 1.5 L of pelletizedfluorinated Cr₂O₃ catalyst. Nickel mesh was placed at the top and at thebottom of reactor to support the catalyst. A multi-point thermocouplewas inserted at the center of the reactor. A feed containing about 10.0wt % 1233zd(E) and 86.3 wt % 1233zd(Z) was introduced into the reactorat the rate of 0.7 lb/hr. The feed was vaporized prior to entering thereactor preheater. The reactor temperature for this experiment wasvaried between 100° C. and 200° C. The temperature gradient throughoutthe reactor never exceeded 3-5° C. Samples of reaction products weretaken every hour and GC analysis of those samples is given in Table 4.

TABLE 4 Reaction Temp. Area Percent by GC ° C. 1233zd(E) 1233zd(Z)Others Initial 10.0 86.3 3.7 103 69.6 27.9 2.5 104 69.8 27.9 2.4 12870.2 27.6 2.2 128 65.0 32.8 2.2 128 62.8 35.0 2.2 128 60.9 36.9 2.2 15160.8 37.1 2.1 151 61.8 36.2 2.0 151 62.4 35.6 2.0 151 58.9 39.0 2.1 18162.2 35.8 2.0 199 68.3 29.4 2.3

Example 10

This example illustrates a continuous vapor phase fluorination reactionof 1,1,1,3,3-pentachloropropane(HCC-240fa)+3HF→1-chloro-3,3,3-trifluoropropene (1233zd)+4HCl. Thefluorination catalyst for the experiment was fluorinated Cr₂O₃.

A continuous vapor phase fluorination reaction system including N₂, HF,and organic feed systems, feed vaporizer, superheater, 2″ ID monelreactor, acid scrubber, dryer, and product collection system was used tostudy the reaction. The reactor wad loaded with 2135 grams offluorinated Cr₂O₃ catalyst which is about 1.44 liters of catalyst. Thereactor was then heated to a reaction temperature of about 275° C. witha N₂ purge over the catalyst after the reactor had been installed in aconstant temperature sand bath. The reactor was maintained at about 2psig of pressure. HF feed was introduced to the reactor (via thevaporizer and superheater) as a co-feed with N₂ for 15 minutes when theN₂ flow was stopped. The HF flow rate was adjusted to 1.0 lb/hr and then1,1,1,3,3-pentachloropropane (HCC-240fa) feed was introduced in to thereactor (via the vaporizer and superheater). The feed rate of HCC-240fawas kept steady at about 1.2 lb/hr and HF feed was kept steady at 1.0lb/hr for about a 9 to 1 mole ratio of HF to 240fa. Once the reactionstarted, the catalyst bed temperature was adjusted to about 328 to about332° C. The average composition of the material at the exit of thereactor was about 83.0 GC area % HCFO-1233zd(E), 8.95 GC area %HCFO-1233zd(Z), 3.48 GC area % 1234ze(E), 2.06 GC area % 245fa, 1.41 GCarea % 1234ze(Z), and 0.08 GC area % 3,3,3-trifluoropropyne. Duringabout 200 hours on stream, the position of a hot spot inside thecatalyst bed moved from the inlet to the exit section of the reactorindicating partial deactivation of the catalyst, but the conversion of240fa was remained at 100% throughout the run.

Example 11

The fluorinated Cr₂O₃ catalyst deactivated after 200 hours of on-streamtime as described in Example 10 was regenerated by the followingprocedure:

The reactor was heated to 300° C. while flowing N₂ at a rate of 5000cc/min.

Synthetic air was introduced after reactor temperatures stabilized. Airflow was started with a rate that gave 0.5% O₂. Gradually, with 0.25% O₂increments, air flow was increased to achieve O₂ concentration of 2.0%.Then, reactor hot-spot was brought to 360° C. and the air flow rate wasgradually, in 0.5-1.0% increments, increased to achieve O₂ concentrationof 5.0%. Careful adjustments of reactor heater temperature were neededto avoid overheating the reactor above 380° C.

The reactor was maintained at a 360-375° C. hot spot temperature whileflowing 5% O₂/N₂ until the hot spot reached the top of the catalyst bed.Then, without changing reactor heater temperature, O₂ flow wasmaintained until the reactor temperature approached that of the reactorheater. Then, the reactor was purged with N₂ for 5 hours to removeresidual oxygen and moisture. That completed the regeneration of thecatalyst and the reactor was brought to 275° C. to prepare it forre-fluorination with HF.

The 240fa+3HF→1233zd+4HCl reaction was restarted at the same operatingconditions described in Example 10. The position of the hot-spot movedback to the inlet of the reactor. The conversion of 240fa was about100%.

Example 12

This example is similar to Example 10 except that reactor temperaturewas varied between 310° C. and 350° C., reaction pressure was variedbetween 2 psig and 25 psig, 240fa feed rate was kept constant at 1.2lb/hr and HF feed was varied to achieve HF:240fa molar ratio between 6.3and 9. The effects of the reaction conditions are presented in Table 5.

TABLE 5 HF:240fa Mole Pressure Temp. t-1234ze c-1234ze 245fa t-1233zdc-1233zd ratio psig ° C. GC. % GC. % GC. % GC. % GC. % 9.0 25 310 3.09150.7932 6.0226 78.8773 8.8861 9.0 2 310 3.3722 0.9808 3.3337 80.89248.6707 9.0 25 350 4.1646 1.0482 3.1154 80.4220 9.3069 9.0 2 350 3.98631.1476 1.4706 82.0814 9.7896 6.3 2 350 2.1189 0.6183 0.7510 85.15459.4874 6.3 25 310 1.7233 0.4568 2.6915 84.3689 8.9174 6.3 2 310 1.86760.5473 1.4649 85.6759 9.0857 6.3 25 350 2.2432 0.6582 1.6793 83.96249.5061

Example 13

This example illustrates purification of the target productHCFO-1233zd(E).

A distillation column was charged with 118.9 lb of the acid free crudeHCFO-1233zd product. The composition of the crude mixture is shown inTable 6.

TABLE 6 Composition of crude HCFO-1233zd product charged in to thedistillation column Component Conc. (GC %) CF₃CCH 3.74 1234ze(E) 7.651234ze(Z) 1.06 245fa 1.75 1233xf 0.11 1233zd(E) 81.38 1233zd(Z) 4.08others 0.23

The distillation column consisted of a 10-gallon reboiler, 2-inch ID by10 feet long column packed with Propack high efficiency distillationcolumn packing, and a shell and tube condenser. The column had about 30theoretical plates. The distillation column was equipped withtemperature, pressure, and differential pressure transmitters. Thedistillation was run at a pressure of about 40-50 psig during the lightscut and at a pressure of about 30 psig during the main, HCFO-1233zd(E),cut. The distillate was sampled and analyzed by GC at regular intervals.Two separate cuts were collected: lights cut and main cut. Hiboilerswere drained from the bottom of the reboiler after the distillation wascomplete. The recovery of essentially pure HCFO-1233zd(E) was about 76%.The compositions and weights of three cuts are listed in Table 7.

TABLE 7 Composition of the distillation cuts collected during thedistillation described in the Example 13. 1233zd(E) cut LIGHTS BOTTOMS73.8 lb 38.3 lb 6.8 lb Concentration Concentration ConcentrationComponent (GC %) (GC %) (GC %) CF₃CCH — 11.6 — 1234ze(E) — 23.75 —1234ze(Z) — 3.3 — 245fa — 5.44 — 1233xf — 0.35 — 1233zd(E) >99.99 55.2926.18 1233zd(Z) — — 71.29 Others — 0.27 2.53

Example 14

This example illustrates the recovery of anhydrous HF from a mixture ofHF and HCFO-1233zd.

A mixture consisting of about 70 wt. % trans-HCFO1233zd and about 30 wt.% HF is vaporized and fed to the bottom of a packed column at a feedrate of about 2.9 lbs per hour for about 4 hours. A stream of about 80wt. % sulfuric acid (80/20 H₂SO₄/H₂O) with about 2% HF dissolved thereinis fed continuously to the top of the same packed column at a feed rateof about 5.6 lbs per hour during the same time frame. A gaseous streamexiting the top of the column comprises trans-HCFO1233zd with less than1.0 wt. % HF therein. The concentration of HF in the sulfuric acid inthe column bottoms increases from 2.0 wt. % to about 15 wt. %.

The column bottoms containing sulfuric acid and about 15 wt. % HF iscollected and charged into a 2-gallon Teflon vessel. The mixture isheated to about 140° C. to vaporize and flash off HF product, which iscollected. The collected HF product contains about 6000 ppm water and500 ppm sulfur.

The HF collected from flash distillation is distilled in a distillationcolumn and anhydrous HF is recovered. The recovered anhydrous HFcontains less than 50 ppm of sulfur impurities and less than 100 ppmwater.

Example 15

This example demonstrates the use of the recycle column.

A representative HCFO-1233zd vapor phase reactor effluent mixture asdetermined in Examples 10 and 11 is charged into a batch distillationcolumn. The distillation column includes a 10-gallon reboiler, 2-inch IDby 10 feet long column packed with Propack high efficiency distillationcolumn packing, and a shell and tube condenser with −40° C. coolant flowcapability. The column has about 30 theoretical plates. The distillationcolumn is equipped with temperature, pressure, and differential pressuretransmitters. The distillation column feed mixture is about 30 wt. % HF,37 wt. % HCl and 33% HCFO-1233zd(E) crude. The distillation is run at apressure of about 100 psig and a differential pressure (delta P) of15-20 inches of water. Both the distillate and reboiler are sampledperiodically and analyzed for organic, HF, and HCl using gas and ionchromatography. Initially, HCl, organic and HF are observed in bothsamples. As more material is removed as distillate the concentration ofthe reboiler changes. First, the concentration of HCl decreases until itis undetectable. The distillation proceeds until the concentration oforganic in the reboiler sample decreases to only trace amounts asanalyzed using gas chromatography. The material remaining in thereboiler at the conclusion of the distillation is essentially pure HF.The recovered HF (reboiler bottoms) is then used to demonstrate recycleof recovered HF back to the vapor phase fluorination reactor and workssatisfactorily.

Example 16

This example demonstrates the HF recovery by phase separation.

It is visually observed using a Teflon cylinder that HF and HCFO-1233zdform a heterogeneous mixture. The separation of HCFO-1233zd and HF layeris tested in the temperature range from +10° C. to −30° C.

The phase-separation of a mixture containing HCFO-1233zd and HF isperformed in the temperature range of −30° C. to +10° C. A 500 ml SSsample cylinder is used. The temperature of the cylinder is controlledwith ethanol circulating through the coil wrapped around the cylinder. Athermocouple is attached to the outside wall of the cylinder (betweencooling coil and the cylinder wall) and positioned in the middle of thecylinder to measure the temperature. The cylinder is also equipped withsampling valves at the bottom and the top of the cylinder. To thecylinder is charged 100 g of anhydrous HF and 250 g of HCFO-1233zd(E).The weight ratio HF:HCFO-1233zd(E) is 28.6:71.4. The cylinder is paddedwith nitrogen to 15 psig at −30° C. to allow sampling. Samples are takenfrom the bottom of the cylinder into Tedlar gas sample bags that contain5 grams of distilled water for the purpose of absorbing HF. The firstsample is taken two hours after the cylinder reaches the desiredtemperature. HF concentration is determined by titration with 0.1 N KOHof the aqueous phase of the sample bags. HF concentration in samplestaken after 2 hours at given temperature is presented in Table 8.

HF concentration in the HF layer is analyzed after the organic layer wasremoved from the system. KOH titration showed that concentration of HFin the acid layer was about 70±5%.

TABLE 8 HF concentration in the samples of the bottom (organic) phasetaken after equilibrating the contents of the phase-separator for 2hours at given temperature Temperature HF concentration in bottom (° C.)(organic) phase (wt. %) −30 1.00 −20 1.25 −10 2.75 0 3.25 10 4.00

It should be understood that the foregoing description is onlyillustrative of the present invention. Various alternatives andmodifications can be devised by those skilled in the art withoutdeparting from the invention. Accordingly, the present invention isintended to embrace all such alternatives, modifications and variancesthat fall within the scope of the appended claim.

1. A method for producing a chlorofluoroalkene comprising: providing aliquid reaction admixture comprising hydrogen fluoride, a fluorinatedmetal chloride catalyst and one or more hydrohalocarbons selected fromthe group consisting of 1,1,1,3,3-pentachloropropane,1,1,3,3-tetrachloropropene, 1,3,3,3-tetrachloropropene, and combinationsthereof, wherein said hydrogen fluoride and said one or morehydrohalocarbons are present in a HF:organic molar ratio of greater thanabout 3:1 and wherein said fluorinated metal chloride catalyst isselected from the group consisting of partially or fully fluorinatedTiCl₄, SnCl₄, TaCl₅, SbCl₃, FeCl₃, and AlCl₃; and reacting said hydrogenfluoride and said one or more hydrohalocarbons in the presence of saidcatalyst in a liquid phase and at a reaction temperature of about 85° C.to about 120° C. to produce a reaction product stream comprising(E)1-chloro-3,3,3-trifluoropropene, hydrogen chloride, unreactedhydrogen fluoride, entrained catalyst,(Z)1-chloro-3,3,3-trifluoropropene, and optionally unreactedhydrohalocarbon mixture wherein said product stream has a weight ratioof (E)1-chloro-3,3,3-trifluoropropene to(Z)1-chloro-3,3,3-trifluoropropene of greater than
 1. 2. The method ofclaim 1 wherein said reaction temperature is about 90° C. to about 110°C.
 3. The method of claim 1 wherein said reaction temperature is about95° C. to about 100° C.
 4. The method of claim 1 further comprising:contacting said reaction product stream with (1) a stripping columnequipped with a heat exchanger to produce a first crude product streamcomprising a majority of said hydrogen chloride, a majority of said(E)1-chloro-3,3,3-trifluoropropene, optionally a majority of said(Z)1-chloro-3,3,3-trifluoropropene, and at least a portion of saidunreacted hydrogen fluoride, wherein said portion is, at least, anamount sufficient to form an azeotrope with one or more of said(E)1-chloro-3,3,3-trifluoropropene or said(Z)1-chloro-3,3,3-trifluoropropene, and (2) a reflux componentcomprising a majority of said entrained catalyst, hydrogen fluoride,partially fluorinated intermediates, and unreacted one or morehydrohalocarbons; and returning said reflux component to said reactionadmixture.
 5. The method of claim 4 further comprising: distilling saidreaction product stream to produce a first recycle stream comprisingunreacted hydrogen fluoride from said reaction product stream, unreactedone or more hydrohalocarbons and partially fluorinated intermediatesfrom said reaction product stream, and an overhead stream comprising amajority of said (E)1-chloro-3,3,3-trifluoropropene,(Z)1-chloro-3,3,3-trifluoropropene, hydrogen chloride, and a portion ofsaid hydrogen fluoride from said reaction product stream; and recyclingsaid first recycle stream to said liquid reaction admixture.
 6. Themethod of claim 4 further comprising: distilling said first crudeproduct stream to produce a first by-product stream comprising amajority of said hydrogen chloride from said first crude product stream,and a second crude product stream comprising a majority of said(E)1-chloro-3,3,3-trifluoropropene from said first crude product stream,optionally a majority of said (Z)1-chloro-3,3,3-trifluoropropene fromsaid first crude product stream, and a majority of said unreactedhydrogen fluoride from said first crude product stream.
 7. The method of6 further comprising: separating said second crude product stream into:a recycle stream comprising a majority of said unreacted hydrogenfluoride from said second crude product stream, and a third crudeproduct stream comprising a majority of said(E)1-chloro-3,3,3-trifluoropropene from said second crude productstream, and optionally a majority of said(Z)1-chloro-3,3,3-trifluoropropene from said second crude productstream, and optionally a portion of said unreacted hydrogen fluoridefrom said second crude product stream; and introducing said secondrecycle stream into said liquid reaction admixture.
 8. The method ofclaim 7 wherein said separating comprises: contacting said second crudeproduct stream with a stripping composition comprising sulfuric acid toproduce a sulfuric acid-hydrogen fluoride admixture stream and saidthird crude product stream; and separating at least a portion of thehydrogen fluoride from said admixture to form said second recycle streamand a sulfuric acid recycle stream.
 9. The method of claim 7 whereinsaid separating comprises: cooling said second crude product stream toproduce a composition having a hydrogen fluoride-rich top layer and anorganic-rich bottom layer; separating said top and bottom layers to formsaid second recycle stream, wherein said second recycle stream comprisessaid top layer and said third crude product stream, wherein said thirdproduct stream comprises said bottom layer.
 10. The method of claim 7further comprising: distilling said third crude product stream toproduce a purified product stream comprising a majority of said(E)1-chloro-3,3,3-trifluoropropene and a second by-product streamcomprising said a majority of said (Z)1-chloro-3,3,3-trifluoropropene;and contacting said second by-product stream with a heated surfacemaintained at a temperature of about 50° C. to about 350° C. and in thepresence of an isomerization catalyst to produce a third recycle streamcomprising (E)1-chloro-3,3,3-trifluoropropene isomerized from said(Z)1-chloro-3,3,3-trifluoropropene; introducing said third recyclestream to said third crude product stream.
 11. An integrated system forproducing a hydrofluoroolefin comprising: a. one or more feed streamscumulatively comprising hydrogen fluoride and one or morehydrohalocarbons selected from the group consisting of1,1,1,3,3-pentachloropropane, 1,1,3,3-tetrachloropropene,1,3,3,3-tetrachloropropene, and combinations thereof; b. a liquid phasereactor charged with a liquid phase fluorination catalyst and maintainedat a first temperature of about 85° C. to about 120° C., wherein saidliquid phase reactor is fluidly connected to said one or more feedstreams and wherein said liquid phase fluorination catalyst is selectedfrom the group consisting of partially or fully fluorinated TiCl₄,SnCl₄, TaCl₅, SbCl₃, FeCl₃, or AlCl₃; c. a stripping system comprising astripping column having an average temperature maintained at a secondtemperature of about 10° C. to about 40° C. below said firsttemperature, a reflux stream fluidly connected to said stripping column,and a first crude product stream fluidly connected to said strippingcolumn, wherein said reflux stream is fluidly connected to said liquidphase reactor; d. a hydrogen chloride removal system comprising a firstdistillation column, a hydrogen chloride by-product stream fluidlyconnected to said first distillation column, and a second crude productstream fluidly connected to said first distillation column, wherein saidfirst distillation column is fluidly connected to said stripping column;e. a hydrogen fluoride recovery system comprising a sulfuric acidstripping and recycle system or a phase separation vessel, a secondrecycle stream comprising hydrogen fluoride fluidly connected to saidsulfuric acid stripping and recycle system or a phase separation vessel,a third product stream comprising (E) and (Z)1-chloro-3,3,3-trifluoropropene fluidly connected to said sulfuric acidstripping and recycle system or a phase separation vessel, wherein saidsulfuric acid stripping and recycle system or a phase separation vesselis fluidly connected to said second crude product stream; f. a1-chloro-3,3,3-trifluoropropene purification system comprising a seconddistillation column fluidly connected to said third product stream; afinal product stream comprising (E) 1-chloro-3,3,3-trifluoropropenefluidly connected to said second distillation column; a secondby-product stream fluidly connected to said distillation column, aisomerization reactor fluidly connected to said second by-productstream; an a product recycle stream fluidly connected to saidisomerization reactor and said second distillation column.
 12. A processfor preparing (E)1-chloro-3,3,3-trifluoropropene comprising: providing1,1,1,3,3-pentachloropropane with hydrogen fluoride in a vapor phasereaction mixture and in the presence of a fluorinated catalyst in areactor wherein said hydrogen fluoride and said1,1,1,3,3-pentachloropropane are present in a HF:organic molar ratio ofgreater than about 3:1; and reacting 1,1,1,3,3-pentachloropropane withhydrogen fluoride in the presence of said catalyst in a vapor phase andat a reaction temperature of about 200 to about 450° C. and a pressureof about 0 to about 160 psig to produce a reaction product streamcomprising (E)1-chloro-3,3,3-trifluoropropene, hydrogen chloride,unreacted hydrogen fluoride, unreacted 1,1,1,3,3-pentachloropropane,reaction by-products, and optionally (Z)1-chloro-3,3,3-trifluoropropene.13. The process of claim 12 wherein the fluorinated catalyst is selectedfrom the group consisting of chromium based catalysts, aluminum basedcatalysts, cobalt based catalysts, manganese based catalysts, nickel andiron oxide based catalysts, hydroxide based catalysts, halide basedcatalysts, oxyhalide based catalysts, inorganic salts thereof andmixtures thereof.
 14. The process of claim 12 wherein the fluorinatedcatalyst is selected from the group consisting of Cr₂O₃, Cr₂O₃/Al₂O₃,Cr₂O₃/AlF₃, Cr₂O₃/carbon, CoCl₂/Cr₂O₃/Al₂O₃, NiCl₂/Cr₂O₃/Al₂O₃,CoCl₂/AlF₃, NiCl₂/AlF₃ and mixtures thereof.
 15. The process of claim 12wherein the fluorinated catalyst is selected from the group consistingof FeCl₃/C, SnCl₄/C, TaCl₅/C, SbCl₃/C, AlCl₃/C, and AlF₃/C.
 16. Theprocess of claim 12 further comprising separating the reaction productstream to produce: a first overhead stream comprising HCl; and a firstbottoms stream comprising unreacted hydrogen fluoride, unreacted1,1,1,3,3-pentachloropropane, (E)1-chloro-3,3,3-trifluoropropene, andoptionally (Z)1-chloro-3,3,3-trifluoropropene;
 17. The process of claim16 further comprising further separating the first bottoms stream toproduce: a recycle stream of HF and unreacted1,1,1,3,3-pentachloropropane, which is optionally recycled back to thevapor phase reaction; and a stream that includes reaction by-products,(E)1-chloro-3,3,3-trifluoropropene, and optionally(Z)1-chloro-3,3,3-trifluoropropene;
 18. The process of claim 17comprising further separating said stream to produce: a second overheadstream of reaction by-products with a lower boiling point than(E)1-chloro-3,3,3-trifluoropropene; and a second bottoms stream ofreaction by-products with a higher boiling point than(E)1-chloro-3,3,3-trifluoropropene, (E)1-chloro-3,3,3-trifluoropropeneand optionally (Z)1-chloro-3,3,3-trifluoropropene.
 19. The process ofclaim 18 further comprising separating the second bottoms stream toproduce: a third overhead stream comprising(E)1-chloro-3,3,3-trifluoropropene; and a third bottoms streamcomprising reaction by-products with a higher boiling point than(E)1-chloro-3,3,3-trifluoropropene, some(E)1-chloro-3,3,3-trifluoropropene and optionally(Z)1-chloro-3,3,3-trifluoropropene
 20. The process of claim 19 furthercomprising isomerizing (Z)1-chloro-3,3,3-trifluoropropene to produce(E)1-chloro-3,3,3-trifluoropropene.
 21. An integrated system forproducing a hydrofluoroolefin comprising: a. one or more feed streamscumulatively comprising hydrogen fluoride and1,1,1,3,3-pentachloropropane; b. a vapor phase reactor with a vaporphase fluorination catalyst and maintained at a first temperature ofabout 200 to about 450° C. and a pressure of about 0 to about 160 psig,wherein said vapor phase reactor is fluidly connected to said one ormore feed streams; c. a hydrogen chloride removal system comprising afirst distillation column, a hydrogen chloride by-product stream fluidlyconnected to said first distillation column, and a crude product streamfluidly connected to said first distillation column wherein said firstdistillation column is fluidly connected to said vapor phase reactor; d.a hydrogen fluoride recovery system comprising a sulfuric acid strippingand recycle system or a phase separation vessel, a recycle streamcomprising hydrogen fluoride fluidly connected to said sulfuric acidstripping and recycle system or a phase separation vessel, a productstream comprising (E) and (Z) 1-chloro-3,3,3-trifluoropropene fluidlyconnected to said sulfuric acid stripping and recycle system or a phaseseparation vessel, wherein said sulfuric acid stripping and recyclesystem or a phase separation vessel is fluidly connected to said crudeproduct stream; e. a 1-chloro-3,3,3-trifluoropropene purification systemcomprising a second distillation column fluidly connected to saidproduct stream; a final product stream comprising (E)1-chloro-3,3,3-trifluoropropene fluidly connected to said seconddistillation column; a second by-product stream fluidly connected tosaid distillation column, a isomerization reactor fluidly connected tosaid second by-product stream; an a product recycle stream fluidlyconnected to said isomerization reactor and said second distillationcolumn.
 22. A process for preparing (E)1-chloro-3,3,3-trifluoropropenecomprising: a) reacting 1,1,1,3,3-pentachloropropane with hydrogenfluoride in a vapor phase in the presence of a fluorination catalyst ina reactor mixture to produce a reaction product comprising HF, HCl,reaction by-products, and (E+Z)1-chloro-3,3,3-trifluoropropene; b)separating HF, HCl and reaction by-products having boiling points lowerthan (E)1-chloro-3,3,3-trifluoropropene from the reaction product; c)isolating (E)1-chloro-3,3,3-trifluoropropene from the reaction product;and d) isomerizing (Z)1-chloro-3,3,3-trifluoropropene from the reactionproduct into (E)1-chloro-3,3,3-trifluoropropene.
 23. The processaccording to claim 22, wherein HCl is separated from the reactionproduct in a HCl recovery system, HF is then separated from the reactionproduct in an HF recovery system, and then the reaction by-products areseparated from the reaction product by distillation.
 24. The processaccording to claim 22, further comprising recycling separated HF intothe reactor mixture to react with the 1,1,1,3,3-pentachloropropane. 25.The process according to claim 22, further comprising recyclingby-products other than (Z)1-chloro-3,3,3-trifluoropropene having boilingpoints higher than (E)1-chloro-3,3,3-trifluoropropene to the reactormixture.
 26. The process according to claim 22, wherein isomerizing the(Z)1-chloro-3,3,3-trifluoropropene comprises recycling the(Z)1-chloro-3,3,3-trifluoropropene to the reactor mixture.
 27. Theprocess according to claim 22, wherein isomerizing the(Z)1-chloro-3,3,3-trifluoropropene comprises contacting the(Z)1-chloro-3,3,3-trifluoropropene with a catalyst.
 28. The processaccording to claim 27, wherein the catalyst is fluorinated chromiumoxide.
 29. The process according to claim 27, further comprising cyclingeffluent from isomerization to the reaction mixture.